Removing carbon dioxide from waste streams through co-generation of carbonate and/or bicarbonate minerals

ABSTRACT

Apparatuses and methods for removing carbon dioxide and other pollutants from a gas stream are provided. The methods include obtaining hydroxide in an aqueous mixture, and mixing the hydroxide with the gas stream to produce carbonate and/or bicarbonate. Some of the apparatuses of the present invention comprise an electrolysis chamber for providing hydroxide and mixing equipment for mixing the hydroxide with a gas stream including carbon dioxide to form an admixture including carbonate and/or bicarbonate.

CROSS-REFERENCES TO RELATED APPLICATIONS

This application claims priority to U.S. Provisional Patent ApplicationSer. No. 60/718,906, entitled “Removing Carbon Dioxide from WasteStreams through Co-Generation of Synthetic Carbonate Minerals” filedSep. 20, 2005; U.S. Provisional Patent Application Ser. No. 60/642,698,filed Jan. 10, 2005; and U.S. Provisional Patent Application Ser. No.60/612,355, filed Sep. 23, 2004. The entire text of each of theabove-referenced disclosures (including the appendices) is specificallyincorporated by reference herein without disclaimer.

BACKGROUND OF THE INVENTION

1. Field of the Invention

The present invention generally relates to the field of removing carbondioxide and, collaterally with that removal, other pollutants from wastestreams. More particularly, the present invention relates to removingcarbon dioxide and other pollutants from waste streams through theabsorption of carbon dioxide and other pollutants from concentratedflue-gas-like streams and then the co-generation of carbonate and/orbicarbonate materials that entrain and neutralize any incidentalpollutants absorbed.

2. Related Art

Considerable domestic and international concern in both private andcommercial sectors has been increasingly focused over the last fourdecades on emissions from industries into the air. In particular,attention has been focused on the greenhouse gases that have theproperty of affecting the retention of solar heat in the atmosphere,producing the “greenhouse effect.” The greenhouse effect occurs whenincoming heat from the sun is trapped in the atmosphere and hydrosphereof the earth, raising the average atmospheric temperature, oceantemperature, and other mean and average temperature measures of planetEarth, up to and including the point of climatic change; the effect isgenerally agreed as an operating effect in the Earth's thermal balance,though the rates, the extent to which man's combustion of materialsaffects it and the extent, direction, and magnitude of the effect aredebated. Despite the degree of debate, all would agree there is abenefit to removing CO₂ (and other chemicals) from point-emissionsources if the cost for doing so were sufficiently small.

Greenhouse gases are predominately made up of carbon dioxide and areproduced by municipal power plants and large-scale industry insite-power-plants, though they are also produced in any normal carboncombustion (such as automobiles, rain-forest clearing, simple burning,etc.), though their most concentrated point-emissions occur atpower-plants across the planet, making reduction or removal from thosefixed sites an attractive point to effect a removal-technology. Becauseenergy production is a primary cause of greenhouse gas emissions,methods such as reducing carbon intensity, improving efficiency, andsequestering carbon from power-plant flue-gas by various means has beenresearched and studied intensively over the last thirty years.

Reducing carbon intensity involves the alternate use of non-carbonenergy sources such as nuclear, hydroelectric, photovoltaic, geothermal,and other sources of electric power to reduce the percentage of powerproduced through exclusive carbon combustion. While each of thesetechniques of power-generation continues to gain in terms of totalenergy production, the projections of world electricity demand areexpected to increase at rates faster than energy production from thesemethods. Therefore, carbon greenhouse gas emissions are expected toincrease despite growth in non-carbon energy sources.

Improving efficiency has generally focused on techniques of improvingthe combustion of carbon through pre-combustion, decarbonization,oxygen-fired combustion, etc. by first decreasing the amount of CO₂produced and then oxidizing all potential pollutants as completely aspossible. Also, the technique increases the amount of energy generatedper carbon dioxide emission released for improved efficiency. Whilestrides in this area have improved combustion efficiency, there islittle more improvement to be extracted from this field of endeavor.

Attempts at sequestration of carbon (in the initial form of gaseous CO₂)have produced many varied techniques, which can be generally classifiedas geologic, terrestrial, or ocean systems. These techniques areprimarily concerned with transporting generated carbon dioxide tophysical sites and injecting the carbon dioxide into geologic, soil, orocean repositories. Each of these sequestering techniques involves largecosts in preparing CO₂ for transport, accomplishing the transport, andperforming the injection into a “carbon bank.” As such, these techniquesare generally not economically feasible and in many cases consume moreenergy than the original carbon produced.

Sequestration can also include several industrial processes whichinclude scrubbing, membranes, lower-cost O₂, and hydrates. However, eachof these technologies suffer due to the capital plant costs raised touneconomic levels, and the effect of CO₂ capture on the cost ofelectricity is prohibitive.

The referenced shortcomings are not intended to be exhaustive, butrather are among many that tend to impair the effectiveness ofpreviously known techniques for removing carbon dioxide from wastestreams; however, those mentioned here are sufficient to demonstratethat the methodologies appearing in the art have not been altogethersatisfactory and that a significant need exists for the techniquesdescribed and claimed in this disclosure.

SUMMARY

The present invention provides methods and apparatuses for removingcarbon dioxide from waste streams, which methods and apparatusesovercome the problems described above. In general, the invention relatesto methods of removing carbon dioxide from a gas stream using ahydroxide compound to form a carbonate and/or bicarbonate. The carbonateand/or bicarbonate can then be employed in any number of uses, or simplydiscarded.

Various embodiments of the present invention provide advantages overcurrent technologies for filtering carbon dioxide from waste streams.Some potential advantages that are realizable by the invention includethe following.

Unlike processes that physically remove the carbon to a remote site, theco-generation at industrial scale of sodium carbonates and/orbicarbonates from sodium chloride and carbon dioxide by synthesisdirectly converts the undesirable carbon dioxide into chemicals at thepoint of power-generation, potentially eliminating transportation coststo a sequestration site.

Unlike other efforts at decarbonation of flue-gas streams that are notamenable to retrofitting, embodiments of the present invention may beretrofitted to existing power-plants, greatly lowering the capital costsnecessary to implement decarbonation processing. Additionally, thedecarbonation processing is scaleable and can be implemented bypilot-to-intermediate-to-full-scale implementation by the addition ofincremental reactor units.

Unlike other processes in the art, the decarbonation process of certainembodiments sequesters carbon-dioxide into economically useful chemicalsand co-incidentally produces useful by-products such as chlorine gas,sodium carbonate, and hydrogen gas. Because the by-products of thedecarbonation process are economically useful, those values offset thecosts of sequestration, and in properly designed systems, potentiallymake the sequestration process profitable in itself.

Due to the co-incidental process of aggressive scrubbing of emittedflue-gases, other undesirable acidic pollutants such as SO_(X), NO_(X),HgO_(X), and others are aggressively scrubbed in the process.Additionally, the scrubbing process can lead to the entrapment and/orentrainment of other gas stream components and/or contaminants in thesodium carbonate (e.g., ash from coal, etc.), thereby removing them fromthe gas stream.

Particular embodiments of the present invention comprise a method ofremoving carbon dioxide from a gas stream comprising: obtaining ahydroxide in an aqueous mixture; admixing the hydroxide with the gasstream to produce carbonate products (defined as products containing thecarbonate group, CO₃), bicarbonate products (defined as productscontaining the bicarbonate group, HCO₃), or a mixture of carbonate andbicarbonate products in an admixture; and separating said carbonateand/or bicarbonate products from the admixture, thereby removing carbondioxide from the gas stream. The hydroxide can be any form of hydroxide,including but not limited to sodium hydroxide, potassium hydroxide,calcium hydroxide, magnesium hydroxide, and aluminum hydroxide. Those ofordinary skill will understand that it is possible to obtain similarchemistry and decarbonation with any number of hydroxides or mixtures ofhydroxides. In some preferred embodiments, the hydroxide is a sodiumhydroxide.

In certain embodiments, the method further comprises process-controllingthe method to produce substantially only carbonate products orsubstantially only bicarbonate products. In other embodiments, themethod further comprises process-controlling the method to produce amixture of carbonate and bicarbonate products, which mixture can becomprised of about X % carbonate and about Y % bicarbonate, with X-Ycombinations being any of the following: 1-99, 2-98, 3-97, 4-96, 5-95,6-94, 7-93, 8-92, 9-91, 10-90, 15-85, 20-80, 25-75, 30-70, 35-65, 40-60,45-55, 50-50, 55-45, 60-40, 65-35, 70-30, 75-25, 80-20, 85-15, 90-10,91-9, 92-8, 93-7, 94-6, 95-5, 96-4, 97-3, 98-2, or 99-1.

In certain embodiments, the admixing occurs in two separate chambers,with one chamber being used to produce carbonate products and the otherchamber being used to produce bicarbonate products. In otherembodiments, the admixing occurs in a bubble column or series of bubblecolumns. In still other embodiments, separation of carbonate and/orbicarbonate products from the admixture involves a heated-precipitationseparation process. In some embodiments, the heat for the separationprocess is derived from heat exchange with incoming flue-gases. Theseparated carbonate may be in the form of an aqueous slurry or as asolution of hydroxide, carbonate, and water at various concentrations atthe time of separation, and if so, it can then be dried by any of anumber of methods. In some embodiments, the carbonate need not be dried.For example, a slurry of sodium carbonate can be used in treatment ofhard water. Of course, those of skill will know a wide variety of usesto which the carbonate produced via the methods of the invention can beput, for example, slurry mixtures of sodium bicarbonate and sodiumcarbonate can be slurried to tank-car for use in various forms ofdetergent manufacture, in glass manufacture as a flux, etc., as well asthe previously-mentioned water-treatment uses.

In certain embodiments, the method further comprises transporting thecarbonate products to a remote sequestration site; combining thecarbonate products with acid in a neutralization reaction to generatepure carbon dioxide; and injecting the carbon dioxide into a carbonbank. In other embodiments, other components of the gas stream areneutralized and/or entrained/captured in the carbonate-formationprocess, including SO_(X), NO_(X), and mercuric-containing material.

In some embodiments, obtaining the hydroxide comprises: obtaining asalt; admixing the salt with water, steam, or both to produce asolution; and electrolyzing the solution to produce a hydroxide. Incertain embodiments, the solution is electrolyzed using a voltagegreater than or equal to about 5 volts, while in other embodiments thesolution is electrolyzed using a voltage less than about 5 volts. Insome embodiments, the solution is electrolyzed using a voltage between 1volt and 5 volts, including about 1.5 volts, about 2.0 volts, about 2.5volts, about 3.0 volts, about 3.5 volts, about 4.0 volts, or about 4.5volts, or at any range derivable between any of these points.

In certain embodiments, acid is added to the solution before it iselectrolyzed. The acid can be any form of acid that can provideprotonation to the solution, including but not limited to hydrochloricacid. Those of ordinary skill will understand that it is possible toobtain similar chemistry and electrolysis with any number of acids ormixtures of acids. In some preferred embodiments, the acid ishydrochloric acid. In other embodiments, the amount of acid added to thesolution is based on a determination of the optimum protonation ratethat achieves the lowest energy to produce reactants and the highestenergy to recover from products.

In still other embodiments, the electrolyzing step occurs in anelectrochemical cell having a catholyte side and an anolyte side and thecarbonate and/or bicarbonate products are recycled to the catholyte sideof the electrochemical cell. In other embodiments, the energy requiredby the method is supplemented with waste-heat recovered from the gasstream.

Other embodiments of the invention comprise a method of removing carbondioxide from a gas stream comprising: obtaining sodium hydroxide in anaqueous mixture; admixing the sodium hydroxide with the gas stream toproduce sodium carbonate, sodium bicarbonate, or a mixture of sodiumcarbonate and bicarbonate; and separating said sodium carbonate and/orbicarbonate from the admixture, thereby removing carbon dioxide from thegas stream.

In some embodiments, the method further comprises process-controllingthe method to produce substantially only sodium carbonate orsubstantially only sodium bicarbonate. In other embodiments, the methodfurther comprises process-controlling the method to produce a mixture ofsodium carbonate and bicarbonate, which mixture can be comprised ofabout X % sodium carbonate and about Y % sodium bicarbonate, with X-Ycombinations being any of the following: 1-99, 2-98, 3-97, 4-96, 5-95,6-94, 7-93, 8-92, 9-91, 10-90, 15-85, 20-80, 25-75, 30-70, 35-65, 40-60,45-55, 50-50, 55-45, 60-40, 65-35, 70-30, 75-25, 80-20, 85-15, 90-10,91-9, 92-8, 93-7, 94-6, 95-5, 96-4, 97-3, 98-2, or 99-1.

In certain embodiments, the admixing occurs in two separate chambers,with one chamber being used to produce sodium carbonate and the otherchamber being used to produce sodium bicarbonate. In other embodiments,the admixing occurs in a bubble column or series of bubble columns. Instill other embodiments, separation of sodium carbonate and/orbicarbonate from the admixture involves a heated-precipitationseparation process. In some embodiments, the heat for the separationprocess is derived from heat exchange with incoming flue-gases.

In certain embodiments, the method further comprises: transporting thesodium carbonate to a remote sequestration site; combining the carbonateproducts with acid in a neutralization reaction to generate pure carbondioxide; and injecting the carbon dioxide into a carbon bank.

In some embodiments, obtaining sodium hydroxide comprises: obtainingsodium chloride; admixing the sodium chloride with water, steam, or bothto produce brine; and electrolyzing the brine to produce sodiumhydroxide and chlorine gas. In certain embodiments, the brine iselectrolyzed using a voltage greater than or equal to about 5 volts,while in others is electrolyzed using a voltage less than about 5 volts.In some embodiments, the solution is electrolyzed using a voltagebetween 1 volt and 5 volts, including about 1.5 volts, about 2.0 volts,about 2.5 volts, about 3.0 volts, about 3.5 volts, about 4.0 volts, orabout 4.5 volts, or at any range derivable between any of these points.

In some embodiments, acid is added to the brine before it iselectrolyzed. The acid can be any form of acid that can provideprotonation to the solution, including but not limited to hydrochloricacid. Those of ordinary skill will understand that it is possible toobtain similar chemistry and electrolysis with any number of acids ormixtures of acids. In some preferred embodiments, the acid ishydrochloric acid. In still other embodiments, the amount of acid addedto the brine is based on a determination of the optimum protonation ratethat achieves the lowest energy to produce reactants and the highestenergy to recover from products.

In certain embodiments, the electrolyzing step occurs in anelectrochemical cell having a catholyte side and an anolyte side and thesodium carbonate and/or bicarbonate are recycled to the catholyte sideof the electrochemical cell. In other embodiments, the energy requiredby the method is supplemented with waste-heat recovered from the gasstream. In still other embodiments, the method further comprisescollecting the chlorine gas, while in others hydrogen gas is produced.In some embodiments, the hydrogen gas and the chlorine gas are combustedto form hydrochloric acid, which is added to the brine before it iselectrolyzed. In other embodiments, the hydrogen gas is combusted withatmospheric oxygen or oxygen from stock chemicals to produce water,while in others the methods comprise using the hydrogen gas to produceenergy. In some embodiments, separation of sodium carbonate and/orbicarbonate from the admixture involves a heated-precipitationseparation process in which the heat for the separation process isderived from the energy produced by the hydrogen gas. In others, thehydrogen gas is co-burned with coal to improve coal-fired emissions orit is used in a combustion process for fuel-cell recovery of DCelectricity.

In some embodiments, the gas stream is an exhaust stream from a plant,while in others the plant is a power plant that employs a carbon-basedfuel source. In certain embodiments, the exhaust stream comprises CO₂and H₂O.

Particular embodiments of the present invention also comprise anapparatus comprising: a electrolysis chamber comprising at least onecathode and at least one anode, the chamber adapted to produce hydroxideduring use; mixing equipment operably connected to the electrolysischamber and to a conduit adapted to contain a gas stream during use, themixing equipment adapted to admix hydroxide from the electrolysischamber with the gas stream during use to create an admixture in whichcarbon, sulfur, and/or nitrogen compounds in the gas stream can reactwith the hydroxide; and a separation chamber operably connected to themixing equipment and adapted to separate the admixture into a separategas phase and solid and/or liquid phase.

In some embodiments, the electrolysis chamber comprises a membrane cell,a diaphragm, and/or mercury. In certain embodiments, the mixingequipment is a batch reactor or series of batch reactors, while inothers the mixing chamber is a gas/liquid absorption/reaction device orseries of gas/liquid absorption/reaction devices. In other embodiments,the mixing chamber is a crystallization tower or series ofcrystallization towers, while in others it is a bubble column or seriesof bubble columns.

In certain embodiments, the apparatus further comprises a drying chamberoperably connected to the separation chamber and adapted to removeliquid from the solid and/or liquid phase during use, while in others,the drying chamber is adapted to heat the solid and/or liquid phaseduring use. In still other embodiments, the apparatus is further definedas operably connected to a power plant.

In some embodiments, the electrolysis chamber is adapted to producechlorine gas and sodium hydroxide from sodium chloride and water duringuse. In other embodiments, the mixing equipment is adapted to admixhydroxide from the electrolysis chamber with carbon dioxide from the gasstream during use to produce carbonate and/or bicarbonate products.

In still other embodiments, the present invention comprises a method ofdetermining optimal operating voltage and current of an electrochemicalcell for low-voltage operation with respect to increased area, for agiven V/I characteristic at a given protonation. In other embodiments,the invention comprises a method for determining a lower thermodynamiclimit on operating voltage for a given electrolytic cell employed in theprocess. In certain embodiments, a method of defining ecologicalefficiency (∂CO₂/∂E) generally for devices that accomplish work toremove CO₂ from waste streams is provided, while other embodimentsinclude a method of defining ecological efficiency (∂CO₂/∂E)specifically for devices that employ the present invention in any of itsembodiments. Other embodiments include a method for producing extremelypure hydrogen gas at a low price-indifference point, its cost equalingthe retrievable energy content.

The terms “comprise” (and any form of comprise, such as “comprises” and“comprising”), “have” (and any form of have, such as “has” and“having”), “contain” (and any form of contain, such as “contains” and“containing”), and “include” (and any form of include, such as“includes” and “including”) are open-ended linking verbs. As a result, amethod or apparatus that “comprises,” “has,” “contains,” or “includes”one or more steps or elements possesses those one or more steps orelements, but is not limited to possessing only those one or more stepsor elements. Likewise, an element of a device or method that“comprises,” “has,” “contains,” or “includes” one or more featurespossesses those one or more features, but is not limited to possessingonly those one or more features. The term “using” should be interpretedin the same way. Thus, and by way of example, a step in a method thatincludes “using” certain information means that at least the recitedinformation is used, but does not exclude the possibility that other,unrecited information can be used as well. Furthermore, a structure thatis configured in a certain way must be configured in at least that way,but also may be configured in a way or ways that are not specified.

The terms “a” and “an” are defined as one or more than one unless thisdisclosure explicitly requires otherwise. The term “another” is definedas at least a second or more. The terms “substantially” and “about” aredefined as at least close to (and includes) a given value or state(preferably within 10% of, more preferably within 1% of, and mostpreferably within 0.1% of).

As used herein, the terms “carbonates” or “carbonate products” aregenerally defined as mineral components containing the carbonate group,CO₃. Thus, the terms encompass both carbonate/bicarbonate mixtures andspecies containing solely the carbonate ion. The terms “bicarbonates”and “bicarbonate products” are generally defined as mineral componentscontaining the bicarbonate group, HCO₃. Thus, the terms encompass bothcarbonate/bicarbonate mixtures and species containing solely thebicarbonate ion.

In the formation of bicarbonates and carbonates using some embodimentsof the present invention, the term “ion ratio” refers to the ratio ofsodium ions in the product divided by the number of carbons present inthat product. Hence, a product stream formed of pure bicarbonate(NaHCO₃) may be said to have an “ion ratio” of 1.0 (Na/C), whereas aproduct stream formed of pure carbonate (Na₂CO₃) may be said to have an“ion ratio” of 2.0 (Na/C). By extension, an infinite number ofcontinuous mixtures of carbonate and bicarbonate may be said to have ionratios varying between 1.0 and 2.0.

In some preferred embodiments of the invention, hydrochloric acid isadded to the sodium chloride brine feed of a chlor-alkali electrolysiscell, making the following reaction occur:H₂O+NaCl+aHCl+energy→NaOH+(½+a/2)H₂+(½+a/2)Cl₂In this equation, the term “a” is defined as the “protonation factor”and it represents the ratio of protons (H⁺ ions) to sodium ions (Na⁺ions).

As used herein, the term “sequestration” is used to refer generally totechniques or practices whose partial or whole effect is to remove CO₂from point emissions sources and to store that CO₂ in some form so as toprevent its return to the atmosphere. Use of this term does not excludeany form of the described embodiments from being considered“sequestration” techniques.

As used herein, the term “ecological efficiency” is used synonymouslywith the term “thermodynamic efficiency” and is defined as the amount ofCO₂ sequestered by certain embodiments of the present invention perenergy consumed (represented by the equation “∂CO₂/∂E”). CO₂sequestration is denominated in terms of percent of total plant CO₂;energy consumption is similarly denominated in terms of total plantpower consumption.

As used herein, the terms “low-voltage electrolysis” and “LVE” are usedto refer to electrolysis at voltages below about 5 volts.

Descriptions of well known processing techniques, components, andequipment are omitted so as not to unnecessarily obscure the presentmethods and devices in unnecessary detail. The descriptions of thepresent methods and devices, including those in the appendices, areexemplary and non-limiting. Certain substitutions, modifications,additions and/or rearrangements falling within the scope of the claims,but not explicitly listed in this disclosure, may become apparent tothose or ordinary skill in the art based on this disclosure.

Other features and associated advantages will become apparent withreference to the following detailed description of specific embodimentsin connection with the accompanying drawings.

BRIEF DESCRIPTION OF THE DRAWINGS

The following drawings illustrate by way of example and not limitation.The drawings form part of the present specification and are included tofurther demonstrate certain aspects of the present invention. Theinvention may be better understood by reference to one or more of thesedrawings in combination with the description of illustrative embodimentspresented herein:

FIG. 1 is a process-flow diagram showing certain embodiments of thepresent invention.

FIG. 2A shows an apparatus for observing the primary features of oneembodiment of the decarbonation portion of the present invention.

FIG. 2B shows experimental absorption/conversion results.

FIG. 2C shows experimental absorption/conversion results.

FIG. 2D is a chart showing gas/liquid contact distance (m, depth offluid) necessary to remove 90% CO₂.

FIG. 2E is a chart showing product ion ratio vs. percent CO₂ absorbed ina test reactor.

FIG. 3 is a chart showing the thermal behavior approximated by fluidwithin a reaction chamber as the reaction proceeds for the timeindicated.

FIG. 4 is a chart showing a flooding study of a 5′ column.

FIG. 5 is a chart showing typical voltage/current characteristicoperating lines for various anolyte pH and temperature conditions.

FIG. 6 is block diagram of a system including a reactor where hydrogenis not recovered, according to embodiments of the present invention.

FIG. 7 is a block diagram of a system including a reactor where hydrogenis recovered through fuel-cell DC return, according to embodiments ofthe present invention.

FIG. 8 is a chart showing percent CO₂ absorption in a bubble-column vs.fluid depth vs. gas interfacial velocity at low interfacial velocities.

FIG. 9A is a chart showing the theoretical max CO₂ absorption andexperimental results, according to embodiments of the present invention.

FIG. 9B shows assumptions and calculations for a flue-gas model for aplant incorporating certain embodiments of the present invention.

FIG. 9C shows the decarbonator process load and intermediate causticrequirement for a plant incorporating certain embodiments of the presentinvention.

FIG. 9D shows the electrolysis section load and requirements for a plantincorporating certain embodiments of the present invention.

FIG. 9E shows waste-heat calculations for a plant incorporating certainembodiments of the present invention.

FIG. 9F shows energy balance and ecological efficiency calculations fora plant incorporating certain embodiments of the present invention.

FIG. 10 shows the ecological efficiency of various modeled power plantsincorporating embodiments of the present invention.

FIG. 11 is a chart showing percentage power saved for m² area ofnormalized LVE design.

DESCRIPTION OF ILLUSTRATIVE EMBODIMENTS

The present invention relates to sequestration processes in which carbondioxide is removed from waste streams and converted into carbonateand/or bicarbonate products. Embodiments of the methods and apparatusesof the invention comprise one or more of the following generalcomponents: (1) an aqueous decarbonation process whereby gaseous CO₂ isabsorbed into an aqueous caustic mixture and then reacted with thehydroxide to form carbonate and/or bicarbonate products; (2) aseparation process whereby the carbonate and/or bicarbonate products areseparated from the liquid mixture; (3) a brine electrolysis process forproduction of the sodium hydroxide that is used as the absorbent fluidin the decarbonation process; and (4) generation and use of by-productsfrom the decarbonization and electrolysis processes, including chlorinegas, sodium carbonate and bicarbonate, and hydrogen gas. Each of thesegeneral components is explained in further detail below.

I. Overview of Advantages

Like any other method or apparatus that performs work to accomplish anobjective, many embodiments of the present invention consume some energyto accomplish the absorption of CO₂ and other chemicals from flue-gasstreams and to accomplish the other objectives of embodiments of thepresent invention as described herein. However, one advantage of certainembodiments of the present invention is that they provide ecologicefficiencies that are superior to those of the prior art, as explainedin detail in Examples 5 and 6. As is evident from the data in Examples 5and 6, amplified waste-heat-recovery or non-greenhouse-gas-generatedenergy for powering the process, use of even-lower-voltage electrolysis,and improving electrical return from hydrogen-energy-recovery canfurther improve the ecologic efficiency of the process, up to andsurpassing the point where the process is fully powered by waste-heatrecovery (and the recovery of hydrogen-energy), and absorbs virtually100% of power-plant emitted CO₂.

Additionally, one benefit of employing the extreme chemistry of certainembodiments of the present invention such that they absorb the weak-acidCO₂ is that the process virtually completely absorbs the strong acids,SO_(X) and NO_(X), and, to a lesser extent, mercury. Tests usingSO_(X)/Ar and NO_(X)/Ar in charged-load single-stage decarbonators havedemonstrated 99%+ removal of these components of flue-gas (by “99%+,” itis meant that the presence of either pollutant in a 14 L/min flue-gasprocessing case was not detectable in the product air-stream bygas-chromatograph technique, i.e., they were effectively removed). Incertain embodiments of the present invention, the incidental scrubbingof NO_(X), SO_(X), and mercury compounds can assume greater economicimportance; i.e., by employing embodiments of the present invention,coals that contain large amounts of these compounds can be combusted inthe power plant with, in some embodiments, less resulting pollution thanwith higher-grade coals processed without the benefit of theCO₂/absorption process of certain embodiments of the present invention.

Further, the scale-ability of certain embodiments of the presentinvention can be carried out to extreme gradations; i.e., since, incertain embodiments, the process is electrically controlled, thatelectrical power expense can be virtually scaled to the individualmolecule of absorbent produced at any given time. Also, the ability toaccurately determine the amount of CO₂ absorbed is practical and easy:weigh the carbonate/bicarbonate products formed, measure their ion ratioby assay, perform a calculation to determine moles of CO₂ absorbed, andthe CO₂ absorbed is easily confirmed and measured (a factor that maybenefit certain incentive regimes for the removal of CO₂ and otherchemicals in flue-gas).

Another additional benefit of certain embodiments of the presentinvention that distinguishes them from other CO₂-removal processes isthat in some market conditions, the products are worth considerably morethan the reactants required or the net-power or plant-depreciationcosts. In other words, certain embodiments are industrial methods ofproducing chlor-hydro-carbonate products at a profit, whileaccomplishing considerable removal of CO₂ and incidental pollutants ofconcern. All other competing methods are simply additionalcosts-of-operation.

II. Process-Flow Diagram

FIG. 1 depicts a simplified process-flow diagram illustrating general,exemplary embodiments of the apparatuses and methods of the presentdisclosure. This diagram is offered for illustrative purposes only, andthus it merely depicts specific embodiments of the present invention andis not intended to limit the scope of the claims in any way. As shown inFIG. 1, flue-gas (FG) enters the process at 901, potentially afterinitially exchanging waste-heat with a waste-heat/DC generation system.FG, entering in this example as a 300° C. mixture of gases, is firstdelivered by pipe 902 to FG/Bicarbonate Heat Exchanger 903, in which theFG temperature is reduced in this example to 120-140° C. Similarly, theFG continues through Anolyte/FG Heat Exchanger 904 and Catholyte HeatExchanger 905, lowering its temperature to 95° C., and then throughWater/FG Heat Exchanger 929, lowering its temperature further to 30° C.The FG exiting Water/FG Heat Exchanger 929 is then introduced to avalving arrangement, Flue-Gas Temperature Mix Control 931, in which the30° C. FG can be mixed with 120-140° C. flue-gas, delivered also toFlue-Gas Temperature Mix Control 931 by means of Hot Flue-Gas ProcessPipe 906. FG mixtures between 30-140° C. may then be differentiallyintroduced to the bottom of Carbonation/Absorption Column 907, which canbe a packed or unpacked bubble column, in which the gas is injected orsparged, with the effect that the gas forms bubbles that rise throughthe fluid, collecting at Upper Vent 908. In this embodiment, thepartially or wholly decarbonated fluid is then injected and passedthrough Bicarbonation/Conversion Column 909, bubbles through the fluidin that column, is pulled by a blower further, and is ejected to Vent910.

The fluid used in FG/Bicarbonate Heat Exchanger 903 is sodiumbicarbonate/carbonate and various sulfate, nitrate, mercury, andparticulates and aerosols absorbed from the flue-gas in theAbsorption/Conversion Columns (907 and 909). By its contact with the300° C. incoming FG, this liquid fluid is heated to temperaturessufficient to develop significant water vapor pressure, producing steamwhen the fluid is injected into Tank 911, which is then condensed inCondenser 912, with the resulting distilled water being recycled to H₂OReclaim Tank 913, which is further used, after any needed conditioning,to form brine in Brine Mixer 914. The fluid used in the Anolyte/FG HeatExchanger 904 is brine made from the addition of group-1 and group-2salts (in this example NaCl) to water either supplied from Water MainSupply 915 or partially or wholly supplied from H₂O Reclaim Tank 913.This brine is protonated (acidified) by the addition of HCl in the formof HCl gas absorbed by water, or from stock chemical HCl, all controlledby pH Closed-Loop Controller 916. This fluid circulates through AnolyteSection 917 of Electrolysis Cell 933. Similarly, the fluid used inCatholyte/FG Heat Exchanger 905 is NaOH (aq) that is circulated throughCatholyte Section 918 of Electrolysis Cell 933. When the pH of theCatholyte exceeds minimum pH (as proxy for concentration) value atControl Point 919, concentrated NaOH is delivered to Hydroxide StorageTank 920.

The fluid used in Water/FG Heat Exchanger 929 is from a sufficientlylarge water reservoir at a sufficiently cool temperature to accomplishthe heat-exchange. In some embodiments, this heat exchange system can beused as a “pre-warming” treatment for bicarbonate/carbonate solutionsprior to further heat exchange in Flue-Gas/Bicarbonate Heat Exchanger905. Also, in some embodiments, the tanking for Water Main Supply 915,H₂O HX Storage Tank 937, and H₂O Reclaim Tank 913 may be partially orwholly consolidated.

Protonated brine circulating through Anolyte Section 917 of ElectrolysisCell 933 is acted upon by the process of electrolysis, forming chlorinegas that is collected and moved through Chlorine Gas Line 921 to, inthis example, Sodium Hypochlorite Reactor 924. Sodium ions and hydrogenions (protons) are transported through the membrane of Electrolysis Cell933 into Catholyte Section 918. Here, sodium ions displace hydrogen ionsin the water, allowing the formation of hydrogen gas, which is takenaway in Incoming Pure Hydrogen Gas Piping 922 to H₂/O₂ Fuel Cell 923,where it is combined with atmospheric O₂ to produce DC electricity,which is recycled to Electrolysis Cell 933 in this example, and purewater, which is recycled through Pure Water Recovery Loop 935 to H₂OReclaim Tank 913. The chlorine gas delivered to Sodium HypochloriteReactor 924 is contacted (bubbled) through sodium hydroxide delivered tothe reactor from Hydroxide Storage Tank 920. Sodium hypochlorite resultsand is tanked to market or further use as a feed-stock chemical. Some ofthe chlorine and HCl gas produced (using the super-stoichiometricamount, a, will produce a continuous recycle of HCl excepting make-uplosses) is here combusted in HCl Fuel-cell 925, which is then recycledthrough HCl Acid Gas Reflux Line 926 to Brine Mixer 914.

Hydroxide produced and stored, or made from stock chemicals and tankedin Hydroxide Storage Tank 920, is the absorbent fluid introduced toCarbonation/Absorption Column 907. It is then passed throughBicarbonation/Conversion column 909 and is then delivered (as abicarbonate/carbonate mixture in water) to FG/Bicarbonate Heat Exchanger903. After removing water through evaporation, a product slurry ofconcentrated bicarbonate/carbonate is delivered to Product BicarbonateTank 927, where it can be drawn for further processing or refinement, orit can be sent to disposal or to market.

Each of the generalized components of the apparatuses and methods of thepresent disclosure as described above and shown in FIG. 1 are describedin further detail below.

III. Aqueous Decarbonation (Absorption) of CO₂ from Waste Streams andits Conversion into Carbonate and Bicarbonate

As noted above, in certain embodiments, the apparatuses and methods ofthe present disclosure employ an aqueous decarbonation process wherebygaseous CO₂ is absorbed into an aqueous caustic mixture and then reactedwith the hydroxide to form carbonate and bicarbonate products. In manyembodiments of the present invention, unlike other capture/sequestrationschemes, sodium hydroxide is used as the primary absorbent fluid. Sodiumhydroxide, in various concentrations, is known as a ready absorber ofCO₂. When carbon dioxide is brought into contact with aqueous sodiumhydroxide, a continuum of products that range from pure sodiumbicarbonate (NaHCO₃) to pure sodium carbonate (Na₂CO₃) can be formed,and differing conditions can be produced that will drive the equilibriumeither direction, even unto completion (or its near vicinity) and insufficient concentration (by either process chemistry or removal ofwater by various means) the precipitation of bicarbonate, carbonate, ora mixed precipitate containing both compounds.

When carbon dioxide is brought into contact with aqueous sodiumhydroxide, the fluid within the reaction chamber approximates thebehavior shown in FIG. 3 as the reaction proceeds for the timeindicated. The two temperature-excursion phases correspond and identifytwo distinct reaction regimes:

-   -   (1) An initial absorption phase in which CO₂ is readily        absorbed. The absorption ability of the fluid declines as the OH        concentration declines, and absorption ends and in some        instances reverses when OH ion concentration is consumed. The        reaction is exothermic during this portion and forms almost        exclusively carbonate.    -   (2) A secondary conversion phase in which CO₂ is not-readily        absorbed. The passage of the flue-gas through the mixture does        not cause any net CO₂ absorption by the fluid, but the fluid is        significantly cooled by loss of heats of vaporization due to any        evaporation of water, by any loss of CO₂ to the vapor state, and        by any endothermic reactions taking place. During this phase,        sodium carbonate already formed in solution is converted to        sodium bicarbonate, by the following required net stoichiometry:        Na₂CO₃ (aq)+H₂O (l)+CO₂ (aq)→2NaHCO₃ (aq)

This sequence of carbonation-first and then bicarbonation-second isreproducibly demonstrable by repeatedly running the apparatus in FIG. 2A(as explained in detail in Example 3) to and past the absorption limitsof the fluid with different concentrations of absorbent.

The two phases are distinguished by the characteristics shown in Table 1below.

TABLE 1 Thermo- CO₂ [OH] Phase dynamics Product Absorption PresenceCarbonation Exothermic Na₂CO₃ Robust Plentiful Bicarbonation EndothermicNaHCO₃ Reduces, Nil De minimis or negative

While embodiments of the present invention could use the same chamber toaccomplish these two processes in situ, the different natures of thereactions suggest that separating the reactions into two chambers andoptimizing them separately is the proper path for preferred embodiments.Irrespective of the “internal arrangement” of transfer devices (i.e.,the degree of batch-vs-continuous, few numbers of chambers, vessels,stages, etc.), the fundamental two processes occur in this sequence atmolarities sufficient to provide good absorption.

Therefore, since some embodiments of the present methods and apparatusescan be process-controlled to produce pure or near-pure sodiumbicarbonate, some embodiments of the present invention therefore captureone carbon per sodium produced by electrolysis, instead of ½ (animprovement in delivered ecologic efficiency nominally 2× that ofcarbonate production). Thus, the amount of electrolysis energy andprocessing spent to produce a mole of hydroxide has double the “normal”absorption power when used to form bicarbonate, compared to theabsorption/energy efficiency of carbonate formation.

In various embodiments of the present invention, all forms of thebicarbonate/carbonate concentration spectrum may be produced. Inpreferred embodiments, the concentrations, temperatures, pressures,flow-rates, etc. of the fluids can be manipulated to optimize theproportion of “available” CO₂ absorbed to optimize the formation ofbicarbonate.

Some embodiments of the invention may control pH in the absorbing fluid(OH ion concentration) as a means of controlling the absorption rate ofthe CO₂ and other gases. In some preferred embodiments, increasedconcentration of salts/carbonates can be used to further drive thereaction to bicarbonate formation. Market pricing of products andeconomic factors may allow operation to produce carbonate-rich productfor a period of time, then bicarbonate-rich product for another period,etc., with the plant average of Na/C then forming a measure of itsefficient use of the ionic species created for absorption/conversion.

By separating the two processes into two distinct chambers andtransitioning between chambers at the point of OH exhaustion,temperature stasis/fall, and absorption attenuation, the manner in whicha decarbonation apparatus can be built and optimized is altered. Aperson of skill in the art will understand that batch,pseudo-continuous, continuous, etc. versions of this simple two-stageprocess decarbonation process can be engineered.

Further, intending to accomplish the absorption with the least energypossible, many preferred embodiments of the invention may employbubble-column reactors (packed or unpacked, with/without horizontalfluid flow, with or without horizontal fluid-flow) that by their naturecreate large liquid/gas contact area to aid mass transport, from whichthe overall design benefits by the freedom to utilize stages with shortstage height (3 m or less) that yet achieve 90%+ absorption with littleresistance or head-of-pressure to overcome in the pumping of the fluids,and therefore are designed with wide horizontal area to achieveindustrial scaling (wide shallow pools or the equivalent in vessels),potentially with horizontal movement to accommodate continuousoperation. Some embodiments of the invention may utilize gas-liquidcontactors of many other configurations, as long as those devices attainthe required gas-liquid contact.

FIG. 4 shows a flooding study of a 5′ column, in which resistance isapproximately 0.01 psig, plus 1.52 psig of head to overcome thefluid-depth of 5′. These losses and other compression-costs are expectedto expend less than 1% of the power plant basis, and as such areconsidered de minimis and not calculated in the examples. FIG. 4confirms that extremely low resistances in the fluid-path will result inextremely low-energy compression to the effect that the ecologicefficiency of the device is not impaired by excess energy spent oncompression or gas handling.

The ecological efficiency of embodiments of the present methods andapparatuses is enhanced by doing the least work possible to absorb CO₂,and one factor that detracts from that efficiency is the amount ofcompression, pumping of fluid and air that is required to accomplishthat process. To that end, two high-efficiency absorbers (capable ofremoving 99% of the CO₂ from an incoming flue-gas stream that is 60% CO₂in N₂) are designed to operate with “short stages” that achieve high CO₂absorption rates.

Preferred embodiments of the invention use a wide-area liquid-gastransfer surface (bubble-column, packed or clear, or its equivalent instatic or moving fluid vessels) to accomplish a high-absorption rate ina short height of fluid absorbent, thereby lowering the resistancenecessary to bring the fluids into contact, and this “short stagesdesign” therefore requires that wide, short “pools” or their equivalentin piping, trenches, vessels, etc. be employed to efficiently absorblarge quantities of CO₂.

The decarbonation reactions of the present disclosure are generallyconsidered by the mainstream industry and all its reference literatureto be mass-transfer-limited. In practice, using packed or un-packedcolumns with wide-area gas-liquid contact absorption inbubble-rising-through-fluid methods, the reaction appears to have littlemass-transfer limitations, or said differently, utilizing the presentmethod of bubble-column design for liquid-gas contact appears toovercome the mass-transfer limitations handily: bubbling with zeropacking through a sparger with only 30 cm of gas/liquid contact distancehas been demonstrated to produce instantaneous rates of 98%+ absorptionof CO₂ (see FIGS. 2B and 2C as discussed in Example 3), and overindustrially-significant timeframes of 15-25 minutes the fluid retainsthe ability to average as much as 80%+ absorption. This is hardlyseriously mass-transfer limited, and practical experimentation, evenwith a simple charged load run to extinction, demonstrates the readymass-transfer of this chemisorption.

Three examples of the design of high-absorption CO₂/NaOH absorptionreactors are explained in detail in Examples 1-3. The conclusions drawnfrom Examples 1-3 are that high-absorption rates with short stages ofNaOH is proven and demonstrated as industrially capable of removinghigh-percentages of incoming CO₂ at low-resistance, in vessels ofmanufacture-able dimensions.

In summary, certain embodiments of the present invention, with respectto the decarbonation portion of the methods and apparatuses, compriseone or more of the following attributes:

-   -   (1) use of short-stages to achieve high-absorption rates of CO₂        in a carbonation phase of the reaction;    -   (2) separation and processing of carbonated fluid in a        bicarbonation process through continued contact with CO₂-bearing        process gas (or other CO₂-bearing gas with concentrations of CO₂        greater than the partial pressure of CO₂-reforming from the        absorbent fluid);    -   (3) possession of a process sequence that can be used, by        process control of state variables and concentration, to produce        pure bicarbonate, pure carbonate, and all various mixtures        between; and    -   (4) embodiments of the invention can be as efficient as a 1:1        sodium/carbon absorption ratio; this optimizes the CO₂ absorbed        per kw-hr (a variant of Ecologic Efficiency, (∂CO₂/∂), used in        producing the reactant.        IV. Separation of Products

As noted above, in certain embodiments, the apparatuses and methods ofthe present disclosure employ a separation process by which thecarbonate and bicarbonate products are separated from the liquidsolution. Separation of liquid solution products requires an involvedprocess. The formation of sodium hydrogen carbonate (NaHCO₃ or sodiumbicarbonate) and sodium carbonate (Na₂CO₃ or soda ash) in a liquidequilibrium with sodium hydroxide (NaOH or caustic soda) occurs over awide range of temperatures and pressures and provides differentend-points of the equilibrium given different partial pressures of CO₂.By manipulating the basic concentration, temperature, pressure, reactorsize, fluid depth, and degree of carbonation, precipitates of carbonateand bicarbonate may be caused to occur. Alternatively,carbonate/bicarbonate products may be separated from their water by theexchange of heat with incoming flue-gases, in some preferredembodiments. Further, due to the solubility product constant differencesbetween sodium carbonate and sodium bicarbonate, certain non-intuitiveprocessing points can be reached; e.g., one of the peculiarities of theequilibria of carbonates of sodium in certain caustic solutions is thatthe addition of heat encourages precipitation of solid; also, at certainconditions, carbonates have been demonstrated to self-precipitate fromthe aqueous solution at high (93%+) purity.

Alternatively, in certain embodiments the heat for the separationprocess may be derived from the hydrogen produced in the originalelectrolysis or from creative uses of the waste-heat contained in theincoming flue-gas stream. The crystallization process inherentlypurifies the crystallizing mineral through the well-known process ofpurification by crystallization.

The exit liquid streams, depending upon reactor design, may includewater, NaOH, NaHCO₃, Na₂CO₃, and other dissolved gases in variousequilibria. Dissolved trace emission components such as H₂SO₄, HNO₃, andHg may also be found. In one embodiment, to separate/remove the exitingliquid streams, e.g., removing/separating the water from the carbonates(in this sense of the word, “carbonates” means mixtures of carbonate andbicarbonate, potentially with hydroxides present as well; any separationtechnique applied to any such mixture would likely include adding heatenergy to evaporate water from the mixture), the water may be boiledcausing the water to be evaporated using Reboiler 106, shown in FIG. 6.Alternatively, retaining a partial basic solution (e.g., NaOH atapproximately 1 molal) and subsequently heating the solution in aseparating chamber may cause the relatively pure Na₂CO₃ to precipitateinto a holding tank and the remaining NaOH re-circulates back to Reactor200. In other embodiments, pure carbonate, pure bicarbonate, andmixtures of the two in equilibrium concentrations and/or in a slurry orconcentrated form may then be periodically transported to atruck/tank-car. In other embodiments, the liquid streams may bedisplaced to evaporation tanks/fields, where the liquid, such as water,may be carried off by evaporation.

Referring to FIG. 6, the reactor design shown may recover the energystored in the electrolyzed hydrogen either as a combustion fuel, aboiler gas, or in H₂/O₂ fuel cells. Reactor 200 may be employed toproduce a steady-state operation where NaOH and NaHCO₃ may by producedin approximately 50:50 proportions. The hydrogen gas produced in theoriginal electrolysis may be used to provide heat, and the NaHCO₃ may beprecipitated in Separation Chamber 108 with the remaining NaOH refluxedto Reactor 200. The slurry from Separation Chamber 108 may be providedto Water Treatment Chamber 110, which may be coupled to SeparationChamber 108. Alternatively, the slurry may be stored and subsequentlyprovided to Water Treatment Chamber 110 as needed.

FIG. 7 illustrates another reactor design, according to an embodiment ofthe present invention. In particular, FIG. 7 shows a re-capturing ofsome excess energy used to create the hydrogen by-product. The use of atandem, on-site high-efficiency fuel-cell may allow for a direct-currentrecovery that may be used to supplement and partially supply theelectrolysis current. Mixing Chamber 300 provides an admixtureincluding, but not limited to, a percentage of NaOH, NaHCO₃, and NO_(X),SO_(X), and Hg to Separation Chamber 308. Separation Chamber 308 mayseparate the admixture into solid and/or liquid phases by a providingheat to the admixture. A drying chamber (not shown) of SeparationChamber 308 may remove the liquids from the solid and/or liquid phaseduring the process by providing the heat. A resulting dilute form ofNaOH is provided to Boiler 306 which boils the diluted NaOH into aconcentrated form and provides the concentrate to Mixing Chamber 300 viaa reflux loop. The water from Boiler 306 can be provided to ElectrolysisChamber 302, in particular Brine Mixer 302A. The resultant Na₂CO₃/NaHCO₃(slurry) from Separation Chamber 308 can be provided for commercial use.In one embodiment, the carbonate slurry can be directly or indirectly(e.g., storing the NaHCO₃ for later use in process such as hard-watertreatment) provided to a Water Treatment Plant 310. Alternatively, theNaHCO₃ can be further refined, dried, shipped, and provided for otherindustrial uses.

The release of gaseous products includes a concern whether NaOH orcomponents of same can be released safely, i.e., emitting “basic rain”from a power-plant is equally to be avoided with emitting “acid rain.”However, sodium hydroxide is normally used as a scrubbing element inpower-plant production and is approved for use by the EPA. The handlingof sodium hydroxide in power plants as well as the procedures to avoidbasic release is well-known in the art. For example, a simple andinexpensive condenser/reflux unit may prevent any significant emissionof NaOH in gaseous exhaust.

In a carbonate separation precipitation method according to certainembodiments of the present invention, the carbonate equilibriumsterically binds carbon-dioxide and absorbs the gas on contact, with asubstantially instantaneous conversion to carbonate ion. The reactionchain may be mass-transport limited such that once the carbon-dioxidehas been absorbed by the base, the subsequent ionic reactions occur atrapid pace.

The sodium carbonate equilibrium has a characteristic where astemperature is raised, Na₂CO₃ naturally precipitates and collects, whichmakes it amenable to be withdrawn as a slurry, with some fractional NaOHdrawn off in the slurry. In one embodiment, a bleed-through treatment ofthis slurry with some of the wet chlorine produced in the chlorine cyclemay be used to reduce the NaOH to trace NaCl in the NaHCO₃ at levelsthat either approximate, or are less than, sodium carbonate produced bymining subterranean “trona” or deposits. As such, the sodiumcarbonate/caustic equilibria provides carbon with complete transportfrom gas to liquid to solid. In other embodiments, it may be beneficialto use the carbonate loop as a collection medium to collect a slurry ofash, sodium hydroxide, and other various carbonates and impurities andtruck off the slurry as road-base.

V. Brine Electrolysis for the Production of Absorbent Fluid at LowEnergies

As noted above, in certain embodiments, the apparatuses and methods ofthe present disclosure employ brine electrolysis for production of thesodium hydroxide that is used as the absorbent fluid in thedecarbonation process. Brine electrolysis is an electrochemical processprimarily used in the production of concentrated sodium hydroxide(caustic soda) and chlorine gas, and is typically described throughoutthe relevant literature by the following equation:2NaCl+2H₂O+e−→2NaOH+H₂(g)+Cl₂(g)

Brine electrolysis may be accomplished by three general types ofstandard electrolysis cells: diaphragm, mercury, and membrane cells.Each of these types of cells produces the same output products from thesame input reactants. They differ from each other primarily in the waythe reactants and products are separated from each other.

In one embodiment, a membrane cell may be used due to several factors.First, environmental concerns over mercury have reduced the demand forthe mercury cell. Second, the diaphragm cells may produce a relativelyweak caustic product which contains significant concentrations of saltand chloride ion and requires considerable subsequentreprocessing/separation to remove the significant salt content from thecaustic. Third, improvements in fluorinated polymer technology haveincreased the life-time and electrical efficiency of membrane celltechnology, where lifetimes in excess of 5 years are routinelyguaranteed in the industrial markets. Further, thepower-per-ton-of-caustic efficiencies exceeds those of both diaphragmand mercury cells in preferred implementations.

Membrane cell processing units are typified, but not limited to, thefollowing generalized inputs and outputs:Anode: 26% NaCl input+2275 kwh/ton Cl₂→Cl₂(g)+24% NaOHCathode: H₂O input+e−→30-33% NaOH+H₂(g)

It is noted that the power requirements (e.g., 2275 kwh/ton of chlorine)may depend upon individual electrolysis cell designs. As such, therequirements may vary.

Many preferred embodiments may employ membrane cells in this function.Membrane cells have several advantages over other brine-electrolysisprocesses. First, membrane cells neither contain nor produce anyenvironmentally sensitive emissions (e.g., mercury) and are electricallyefficient when compared with diaphragm and mercury cells. They alsoemploy a concentrated/dilute/make-up NaCl loop such that they may bewell-suited for use as a continuous “salt loop” processing unit. Next,NaOH produced in membrane cells without furtherevaporation/concentration may be a naturally appropriate level ofconcentration for use in a decarbonation process (e.g., 30-33% NaOH byweight). Further, hydrogen produced by membrane cells is “clean,”approximately “electronic grade,” and relatively clear of NaCl or othercontamination. As such, hydrogen may be compressed and tanked off aselectronic-grade H₂ gas, used for power-production on-site such ascombustion mix with low-grade coal or for combustion-technology gains.Alternatively, the hydrogen may be used for a boiler fuel for theseparation processes, which may occur after decarbonation. Membrane celltechnology may also be easily scaled from laboratory to plant-sizeproduction by the addition of small incremental units. Additionally,chlorine gas produced by the membrane process is less “wet” than thatproduced by other standard electrolytic processes. As such, a one-stagecompression cycle may be sufficient for production of water-treatmentgrade chlorine.

The above represents the published and practiced state of the art asnormally accomplished toward the end of producing commercial chlorinegas and caustic soda. However, the aims of certain embodiments of thepresent invention are different in several respects, which leads todifferent chemical techniques being employed to accomplish the differingends of certain embodiments of the present invention.

A. Use of Low-Voltage Electrolysis Techniques

In some embodiments of the present invention, the brine electrolysisincorporates low-voltage electrolysis (LVE) techniques, therebyimproving the thermodynamic efficiency of the process. Certainembodiments of the present invention do not exclusively manufacturecaustic soda as an end-product, but instead use NaOH primarily as anintermediate absorbent to absorb CO₂ from flue-gas prior to its exit tothe environment. Because the chlor-alkali business generally involvesshipping caustic soda from the manufacturing plant to its point-of-use,transporting large quantities of water in the caustic soda isuneconomic, so caustic is generally concentrated to approximately 55 wt% for shipping as a liquid by the energy-intensive removal of water bysteam-evaporation, and in some cases, it is concentrated to an anhydroussolid, generally pelletized form. This concentration is achieved mostgenerally by running electrolysis cells at voltages over 5V (in theparlance, “at significant over-voltage”), which achieves theabove-mentioned 30-35% NaOH, which is then followed by asteam-evaporation cycle to achieve the 55% (or to completely dry theproduct to anhydrous, palletized states, etc.).

Most embodiments of the present invention do not transport NaOH as aproduct, and it is producible at usable concentrations (as an aqueousabsorbent fluid) at voltages significantly below the 5V+ standards usedin chlor-alkali plant operation. The impact of employing LVE on thethermodynamic efficiency of the present invention cannot beover-estimated, since power consumed is exactly to the simple DCequation:P=V(voltage)×I(current)×(current Efficiency)

Because the current is fixed by the electrochemical process (one pair ofelectrons for each molecule, etc.), the power is nearly entirelyregulated by the voltage required to electrolyze (an additional strongfactor is the current efficiency, which is also a function of appliedvoltage). Since embodiments of the present invention employ LVE withvoltages easily demonstrated in standard electrolysis cells, byalteration of operating conditions, as low as 2.1V, through process andgeometric modifications to a standard electrochemical membrane cell,embodiments of the present invention consume less power (kw-hr) for eachamount of NaOH formed than traditional high-voltage chlor-alkalielectrolysis. For these reasons, preferred embodiments of the presentinvention include electrolysis units that are designed to employ allavailable techniques to accomplish low-voltage operation, including, butnot limited to: narrow-gap operation, higher-pressure, wide-areamembranes, altered concentrations of anolyte and catholyte andprotonation-ion ratios. Further aspects of LVE are explained in detailin Example 4.

B. “Point of Indifference” Hydrogen Use

Typical chlor-alkali operation produces hydrogen gas that is eitherburned on site as boiler-fuel for steam-evaporation (see above), or inmany cases, is suppressed entirely by the use of an air-electrode; e.g.,a device that blows air containing oxygen on the cathode causing theimmediate reduction reaction to occur:H₂(g)+½O₂(g)→H₂O(l/g),which alters the electrochemical summation of energies and lowers thevoltage necessary to produce caustic and chlorine (in some cases as lowas 2.8V in various industry literature) at the expense of not producinghydrogen. A benefit to preferred embodiments of the present invention isthat they have no need for hydrogen as a steam-evaporation boiler fuel.

Some embodiments of the present invention may include an “air-electrode”process that suppresses hydrogen production (and zeroes the ability torecover energy from that hydrogen), but many preferred embodimentsutilize the production of hydrogen for either energy return or for useas a chemical feed-stock for other processes.

Plants employing certain embodiments of the present invention thereforehave a “point of indifference” regarding the use of hydrogen soproduced; i.e., the plant produces hydrogen for its energy return (e.g.,60% of its energy content in an example case, for instance), and thatamount of energy has a certain economic value, and instead of using thehydrogen for energy return, the hydrogen may be traded or sold foramounts in excess or equal to that economic value of the energy alone.In this manner, plants of this type represent “hydrogen wells” of asort. Hydrogen is produced and consumed for its electric value, andanyone desiring the hydrogen for other purposes need only trade thatelectric value or its economic equivalent to maintain the economics ofthe process. This embodiment has beneficial implications for future“hydrogen economy” schemes: having a ready source of hydrogen, availableat a price-indifference cost point lower than its innate energy content,may prove a useful and advantageous feature of this technology.

C. Use of Protonated Brine Optimization Techniques

Certain embodiments of the present invention intentionally optimize theproduction of hydrogen instead of suppressing it, by the following meansof Protonated Brine Optimization. The general chlor-alkali industry has,in some occasions, employed the addition of HCl in the form ofhydrochloric acid to the brine (NaCl) cell called the anolyte chamber.This “protonation” of the brine has the effect of lowering the voltagerequired to produce a specified amount of caustic in that it lowers thepower required to make the reactant/absorbent of certain embodiments ofthe present invention. In the normal body of work of chlor-alkaliplants, the cause of this voltage-lowering effect is attributed to“neutralizing NaOH that returns across the membrane,” a process that hasbeen studied versus membrane selectivity, etc. and is well-understood.Generally speaking, this “neutralization” is seen to produce only saltwater [NaOH (returning)+HCl (anolyte)→NaCl+H₂O] and is not shown toalter the stoichiometry of the products. However, one point that is notwell-understood or exploited in the art is that the addition of HClfundamentally does indeed alter the stoichiometry of the equation thus:NaCl+aHCl+H₂O→NaOH+(½+a/2)H₂+(½+a/2)Cl₂

This additional quantity of hydrogen and chlorine produced isparticularly of interest to the present invention, as embodiments of thepresent invention rely upon energy-recovery from hydrogen (which isenhanced when more hydrogen is produced and is more economic when morechlorine gas is produced, per mole of absorbent NaOH manufactured). Thatthis additional bounty of hydrogen-for-energy-recovery andchlorine-for-sale is produced while the overall electrical voltage(therefore power and cost) is reduced leads certain embodiments of thepresent invention to optimize the value “a” for the lowest energy toproduce reactants and the highest energy to recover from products. Theoptimum generally lies between 0.5 and 1.0M HCl in the solution of NaCl(aq). However, the optimization is specific for each cell design,geometry, concentration, temperature, and pressure regime. However, foreach electrochemical cell, there is an optimum protonation rate (“a”)that achieves a lowest-energy operation. In practice, extremely high “a”values (>0.15M at 90° C. for instance) will possibly blister mostcommercially-available membranes in short order. Of course, while HCl isa preferred acid for protonation of brine in the invention, many otheracids, as known to those of skill in the art, can be used to protonatethe brine.

D. Self-Producing the HCl that is Used for Protonation

Because certain embodiments of the present invention both use input HCland produce H₂ and Cl₂ gas, the protonation of brine can be madeself-reciprocating; i.e., H₂ and Cl₂ product gases can be burnt (inefficient fuel-cells or in plain burners) to produce HCl gas, which canthen be recycled to the anolyte for protonation of the brine. The energyrecovered from H₂/Cl₂ combustion is higher than that recovered fromH₂/O₂ combustion. This adds to the thermodynamic efficiency of theinvention.

Some preferred embodiments of the present invention use anabsorption/conversion/regeneration pathway as a “front-endconcentration/absorption process” that is then used to absorb andconcentrate CO₂ for use in geologic, oceanic, or terrestrialsequestration techniques (for example, those that inject CO₂ into carbonbanks) by the following means or similar:

-   -   (1) All hydrogen is combusted to produce HCl gas.    -   (2) All HCl gas is reacted with the sodium bicarbonate so        produced    -   (3) Through the associated neutralization reaction, nearly 100%        pure CO₂ is released, and salt water is regenerated that can be        recycled for further absorption cycles.    -   (4) By this process, the invention is used to absorb, convert,        and then release the CO₂, with the net effect that the gas is        removed from the flue-stream, and concentrated for further        sequestration techniques to then process.

E. Mixing Carbonates and Bicarbonates Back into the Catholytic Fluids

Unlike the chlor-alkali industrial use of chlorine/hydrogen cells, someembodiments of the present invention also recycle carbonate andbicarbonate mixtures to the catholyte (caustic) side of theelectrochemical cell. Such techniques are wide and varied, but eachoperating point for the entire process has an optimum recycle ofbi/carbonate mixtures to the catholyte, as this can, in someconcentrations and conditions, lower cell voltage, and therefore cellpower.

F. Use of Waste-Heat-Recovery for Heating of Fluids

Because certain embodiments of the present invention are employed in thepresence of a power-plant or large emission of CO₂ in the form offlue-gas or other hot gases from combustion, there is ample opportunityfor heat-exchange, unlike standard chlor-alkali processes. For instance,a typical incoming flue-gas temperature (after electro-staticprecipitation treatment, for instance) might well be 300° C. Coolingthat flue-gas to a point less than 300° C., while warming the anolyteand catholyte fluids (which, for LVE, should generally be retained >90°C.) allows certain embodiments of the present invention to operatewithout the power-losses associated with anolyte and catholyte heaters.

G. Use of Waste-Heat-Recovery for Powering Process Equipment

Generally, since the flue-gas that is available at power-plant exits attemperatures between 100° C. (scrubbed typical), 300° C. (afterprecipitation processing), and 900° C. (precipitation entrance), orother such figures, considerable waste-heat processing can be extractedby cooling the incoming flue-gas through heat-exchange with apower-recovery cycle, of which an example is an ammonia-water cycle(“Kalina” patent process, for example), a steam cycle, or any such cyclethat accomplishes the same thermodynamic means. Since certainembodiments of the present invention rely upon DC power to accomplishthe manufacture of the reagent/absorbent for the present invention, theprocess can be directly powered, partially or wholly, by waste-heatrecovery that is accomplished without the normal transformer lossesassociated with converting that DC power to AC power for other uses.Further, through the use of waste-heat-to-work engines, significantefficiencies can be accomplished without an electricity generation stepbeing employed at all. In some conditions, these waste-heat recoveryenergy quantities may be found to entirely power embodiments of thepresent invention.

VI. Generation and Use of By-Products from the Decarbonization andElectrolysis Processes

As noted above, certain embodiments of the apparatuses and methods ofthe present disclosure produce a number of useful by-products from thedecarbonization and electrolysis processes, including chlorine gas,sodium bicarbonate, and hydrogen gas. In some embodiments, the hydrogengas produced by the embodiments of the present invention is incorporatedinto a hydrogen energy recapture loop. In one embodiment, the presentinvention may include current fluorinated polymer membrane technologiesto reduce chlorine-ion migration for the decarbonization process. Theprocess may therefore function without extensive energy and costexpenditure to separate out the chlorine ion; the decarbonation andseparation loops are relatively chloride-free. As such, an embodiment ofa decarbonation reactor of the present invention may use common salt,water, and carbon dioxide exhaust combined with electricity to formchlorine gas, sodium bicarbonate, and power-recovery through hydrogencombustion as follows:2NaCl+H₂O+2CO₂ +e−→2NaHCO₃+Cl₂+[½H₂+atm O₂ →e−]

A. Hydrogen Energy Recapture Loop

Two techniques have developed that may use the hydrogen energy capturedby embodiments of the present invention. The first is co-burning ofhydrogen with coal to improve coal-fired emissions, and the secondtechnique involves hydrogen/oxygen fuel cell recovery of DC electricity.Approximately 22% to 32% of the electrolysis-spent energy may berecovered by the H₂/O₂ combustion. Alternatively, H₂ and Cl₂ may combustinto Cl₂ and HCl directly or through fuel-cell or DC recovery. Sourcesof heat from waste-heat removal, from either the electrolysis reaction,which produces approximately 135° C. caustic soda to tank, or from thedecarbonation process itself, which absorbs various heats-of-solution,heats-of-vaporization, and heat liberated in the exothermic reaction,can be utilized by well-known techniques (i.e., re-boilers, etc.) atpower plant locations for the pre-heating of combustion gases or otheruses.

In one embodiment, commercial fuel-cell production of DC electricity maybe advantageous due to the easy-to-handle and safe operations atsub-atmospheric pressures. Immediate consumption of the producedhydrogen may also directly reduce the electrical load cost for the brineelectrolysis. Further, since the hydrogen-energy recovery cycle may beproduced with off-peak electrical production of H₂, Cl₂, and NaOH, whereH₂ may be subsequently used to provide electricity during on-peak loads,the present disclosure provides for making reactants at low-cost whilesubsequently producing auxiliary high-cost on-peak electricity andsimultaneously performing a decarbonation process. The economic utilityof an H₂ energy recovery cycle to increase the peak power production ofa plant by augmenting the current production with H₂ combustion capacityas either fuel or in a fuel cell may provide for the utility of aself-consumption basis.

Alternatively, given a clean carbon-produced source of electronic gradehydrogen gas, municipal utilities, industrial companies, andpower-generation facilities may prosper from using hydrogen produced inmunicipal bus fleets, trains, and other public or private uses ofhydrogen fuel.

B. Other Uses of By-Products from the Decarbonization Process

In one embodiment, the chlorine gas may be a primary oxidizing reagentused to kill bacteria in water-treatment plants around the world.Chlorine, and the some 100+ derivative feed-stock chemicals that derivefrom it, are often cited as being incorporated in as much as 30%+ of USdomestic GDP. It may also be used in the manufacturing of the largestindustrial chemical, hydrochloric acid. Chlorine is also extensivelyused in the plastics industry as its most common non-carbon reactant.

Sodium carbonate is a by-product of the process described in the presentinvention that is commonly used in the production of soap, detergents,and shampoos world-wide, as well as a flux in the manufacture of glass.Further, power utilities, as well as private homes, utilize sodiumbicarbonate to soften hard water by the following general reaction:CaCl₂(aq)+NaHCO₃(aq)→CaCO₃(precipitate)+NaCl(aq)+HCl(aq)

A similar process may be employed using sodium carbonates andbicarbonates from this process to perform ion-exchanges with multiplegroup 1 and group 2 salts to precipitate various carbonates.

Another example of the by-products produced from the decarbonationprocess is hydrogen gas. Hydrogen gas, in electronic grade, is ahigh-form of energy carrier. Conversely, the hydrogen fuel produced canbe burned with “dirtier” levels of coal to reduce that fuel's emissions,can be burned as boiler feed in the separation process, or can beutilized in hydrogen-vehicle fleets.

In further embodiments, it is possible to accomplish the effect oftransporting carbon dioxide to remote sequestration sites by thefollowing method, or its equivalent:

-   -   (1) At the power-plant site, CO₂ and other pollutants are        absorbed by the process or any of its variants, along with the        production of hydrogen, chlorine, and carbonates.    -   (2) At the remote sequestration site, hydrochloric acid and        carbonates, are again combined in a neutralization reaction that        generates virtually pure CO₂. CO₂ produced by these means may        then be injected into the carbon bank.

By these means, the same net effect of transporting CO₂ between thepower site and the sequestration site is achieved without the physicaltransport of CO₂ by liquification and transport by pipe-line, trucking,etc.

EXAMPLES

The following examples are included to demonstrate preferred embodimentsof the invention. These embodiments provide a feasible economic solutionto reduce or even substantially eliminate the carbon dioxide and otherpollutants from power plants by providing and using the resultingreactants in commercial or industrial processes and/or by recyclingenergy.

It should be appreciated by those of skill in the art that thetechniques disclosed in the examples which follow represent techniquesdiscovered by the inventor to function well in the practice of theinvention, and thus can be considered to constitute preferred modes forits practice. However, those of skill in the art should, in light of thepresent disclosure, appreciate that many changes can be made in thespecific embodiments which are disclosed and still obtain a like orsimilar result without departing from the spirit and scope of theinvention.

Example 1 CO₂NaOH Bubble Column Reactor Design by Graphical Methods

In the bubble-column reactor designed for this example, there are fourprimary flow streams, namely:

-   (1) Liquid flowing into the fluid of the bubble column at a given    volumetric flow-rate (V1=cubic volume of fluid per time); in the    chosen case, the incoming volumetric flow-rate equals the outgoing    volumetric flow rate. Hence, both are VI). In this example, V1=0.001    m³/sec.    -   (2) Vg0=the incoming volumetric flow rate of gas, which will be        partially or wholly absorbed by the absorbent fluid. In this        example, Vg0=0.05 m³/sec.    -   (3) Vg=the exiting volumetric flow-rate of gas. In this example,        Vg=0.02 m³/sec.

A bubble-column reactor was designed as bound by the above conditions.Sixty-percent incoming CO₂ in a flue-gas was to be removed by bubblingthrough a solution of concentrated sodium hydroxide. The reaction ismass-transfer limited. The objective of the example was to calculate thereactor size (height and diameter) necessary to remove 99.9% of thecarbon dioxide. P=2 atm, T=298 K. Using the graphical data available inFIG. 8, this example describes the design of both a tall reactor (2.36m) and a short reactor (0.41 m). FIG. 8 is a chart showing percent CO₂absorption of CO₂ in a bubble-column vs. fluid depth vs. gas interfacialvelocity at low interfacial velocities (Schumpe et al., 1979).

2.36 m Height Column Solution

The conversion approaches 100% at a superficial velocity (Ug0) ofapproximately 0.04 m/s. This velocity is in the bubbly-flow range (inwater-like solutions this is approximately 0.05 m/s). Knowing thevolumetric gas flow requirement (Vg0), the diameter of the column wascalculated:Ac=Vg0/Ug0=0.05/0.04=1.25 m2Dc=2√{square root over (Ac/π)}=2√{square root over (1.25/π)}=1.26 m

Therefore, a conversion of 99% of the incoming 60% CO₂ requires a columnof 2.36 m height, 1.25 m² area, having a diameter of 1.26 m, and a totalvolume of 2.95 m³.

0.41 m Height Column Solution

Conversion in a 0.41 m tall column requires a superficial gas velocityof about 0.02 m/s. Similar to above:Ac=Vg0/Ug0=0.05/0.02=2.50 m2Dc=2√{square root over (Ac/π)}=2√{square root over (2.5/π)}=1.78 m

Therefore, a conversion of 99% of the incoming 60% CO₂ requires a columnof 0.41 m height, 2.50 m² area, having a diameter of 1.78 m, and a totalvolume of 1.03 m³.

In summary, it is demonstrated through this example that shorter columnsare more efficient on a per-volume basis in stripping carbon dioxidefrom a flue gas; for this example, by a factor of 3. Therefore, forpreferred embodiments of the present invention, the design targets shortstages and/or multiple-stage reactors composed of short stages.

Example 2 CO₂/NaOH Bubble-Column Design by Mass-Transfer CoefficientSolution

The objective of this example was to determine the mass-transfercoefficient, kla (moles/sec/volume), from theoretical build-up. It wasdetermined from this example that this correlative method canpotentially lead to inconclusive results; i.e., this example highlightsthe difficulty in predicting actual results from theory due to theindeterminacy of measuring some of the critical parameters. Therefore,only experimental scaling can conclusively determine the result of alarge decarbonation unit.

The following equations for gas-hold-up (εg) and mole-transfer (kLa) arefrom correlations from Akita and Yoshida (1973), and are valid forcarbon dioxide and water systems at relatively large column heights anddiameters (i.e., >0.1 m):

Gas-Hold-Up

$\frac{ɛ\; g}{\left( {1 - {ɛ\; g}} \right)4} = {C*\left\lbrack {g*{Dc}\; 2*\rho\;{L/\sigma}} \right\rbrack{1/8}*\left\lbrack {g*{Dc}\;{2/{vl}}\; 2} \right\rbrack{1/12}*\left\lbrack {\mu\;{g/\sqrt{g}}*{Dc}} \right\rbrack}$andMass-Transfer CoefficientkLa(1/sec)=[Cco2*Dco2−h20/Dc2]*[νL/Dco2−h20]0.2*[g*Dc2*ρL/σ]0.62*[g*Dc3/νl2]*ε11

Where

-   -   εg=gas hold-up factor    -   Cco2=concentration of CO₂ in flue-gas    -   Dc=diameter of column    -   νL=0.0001 m2/sec    -   σ=1 cP=0.1 Pa*sec    -   ρL=998 kg/m3        and since Dco2−h20P=1.661 m2*Pa/sec,        therefore, Dco2−HO=1.661/5.067*10(5)=3.278*10(−6) m2/sec

The driving force is the difference between the equilibriumconcentration of carbon dioxide (Cco2*) and the actual liquid phaseconcentration of carbon dioxide (Cco2), which this example assumes iszero; i.e., the sodium hydroxide present instantaneously neutralizes theaqueous carbon dioxide “acid.” The rate of mole transfer per reactorvolume can therefore be written as:Nco2=kLa*[Cco2*−Cco2]=kLa(Cco2*)Rate of Mole Transfer Needed to Remove 99.99% of the carbon dioxide inthe flue-gas. CO₂ is assumed to be an ideal gas under the conditions ofthe column.

Cubic Volume/Sec:Vg0=0.05 m3/sec∴Vco2=0.6*Vg0=0.03 m3/s

Mole/Sec:Vco2*P/RT=[0.03M3/sec*5 Atm]/[0.082 m3 atm/kmolK*298]*1000 gmmole/1kmole=6.14 mole/secNco2 (mole/sec)=0.999 (moleCO2removed/molesCO2in)*6.14 mole/sec=6.13mole/sec

Other Fluid Properties Necessary for Model Correlation

Setting the initial superficial velocity at the bubbly flow limit(Vg0=0.05 m/s), the area and diameter of the column was calculated:Ac=Vg0/UgO=0.05/0.05=1.0 m2Dc=2√{square root over (Ac/π)}=2√{square root over (1.0/π)}=1.13 m

For the gas-phase hold-up reactions this example sets C=0.2 and assumesthat the superficial velocity (Ug) is the average of the entering andexiting velocities; Ug=0.035 m/s=average (0.05 m/s, 0.02 m/s), and usingan equation solver, it was found that εg=0.055.

Next, the mole transfer rate constant was solved for:

${{kLa}\left( {1\text{/}\sec} \right)} = {\frac{{0.6*3.278*10} - {6\mspace{14mu} m\mspace{14mu} 2\text{/}\sec}}{{\left\lbrack {0.6\mspace{14mu} m} \right\rbrack 2}\;}*\frac{0.0001\mspace{14mu} m\mspace{14mu} 2\text{/}\sec}{{3.278 \times \; 10} - {6\mspace{14mu} m\mspace{11mu} 2\text{/}\sec}}*\left\lbrack \frac{9.8\mspace{14mu}{m/s}\; 2*\left( {0.6\mspace{14mu} m} \right)2*998\mspace{14mu}{kg}\text{/}m\mspace{14mu} 3}{0.0696\mspace{14mu}{kg}\text{/}\sec\mspace{11mu} 2} \right\rbrack*{\quad{{\left\lbrack \frac{\left( {9.8\mspace{14mu} m\text{/}\sec\mspace{14mu} 2} \right)*\left( {0.6\mspace{14mu} m} \right)3}{\left\lbrack {0.0001\mspace{14mu} m\mspace{11mu} 2\text{/}\sec} \right\rbrack 2} \right\rbrack*(0.055)1.1} = {{0.386\text{/}\sec{kLa}} = {0.386\text{/}\sec}}}}}$Going back to the driving force equation, the reactor volume (V) wassolved for:V=Nco2/[(kLa)(Cco2*)]=6.13 mole/sec/(0.386/sec*(103.6 mole/m3))=0.15 m3andHc=0.15 m3/1 m2=0.15M

Therefore, the dimensions of a bubble column are Dc=1.13 m and Hc=0.15m, resulting in a significant difference from actual results in realbubble columns.

Several assumptions can explain the difference in the correlative models(that are considered the best models of this behavior) and the actualresults:

-   -   (1) The sodium hydroxide was assumed to have the properties of        water (density, surface tension, etc.); and    -   (2) The concentration of CO₂ in the solution might not be well        characterized as zero; this is the more likely operand; e.g., if        the CO₂ effective concentration is not zero, then the driving        force is less, and a taller column is needed.

It should also be noted that this theoretical correlation suffers undera condition that is also its strength: because terms such as (νL=0.0001m2/sec) are often squared in the denominator, small variations in thesenumbers produce gigantic effects. This type of theoretical build-up isgood for curve-fitting ex-post-facto, but is not a good predictor ofmass-transfer for design purposes. Further, there are strikingfluid-flow effects on the absorption/conversion rate of CO₂, such thatdesigns at-differing-and-progressive scales are indicated.

Example 3 CO₂NaOH Bubble-Column Design (Depth) from Experimental Data

Note that the reliance of certain embodiments of the present inventionupon the “short stage efficiency theory” described herein (3 m or lessgas-liquid contact distance, or fluid stage height, to achieve >90%absorption) is confirmed by two different calculation techniques thatare consistent with practiced chemical engineering design. However, incertain cases (as noted above) certain simplifying assumptions have beenmade in these design calculations, so experimental verification wasindicated and performed with the Results displayed in FIGS. 2B and 2C(explained in further detail below).

Each of these processing runs obtained a certain CO₂ absorption over acertain gas-liquid contact distance (namely the height of the fluid inunpacked, open bubble-column cases); e.g., 20% absorption through 30 cmof fluid.

If the gas were then passed through a second column of the same designand state-of-conditions, the same absorption would occur again; i.e.,20% of the remaining 80% of initial CO₂ would again be absorbed. Thisrelationship eventually attenuates; however, given the highly absorptivecharacteristics of the absorbent fluid, and the tendency of thechemisorption to persist with robust absorption even of dilute CO₂ inthe flue-gas, this effect is ignored for this example and a design of90% removal is effected.

One can see that further passes of the fluid through the absorptionfluid would again reduce remaining CO₂ by 20%, etc. until sufficientpasses have been accomplished to attain the desired absorption level (inthis case 90%).

This results in a “number of base stages” design that determines thedepth of fluid (in multiple depths of stages of 30 cm height) that isrequired to attain 90%. Assuming each stage absorbs the same %CO₂/distance as the former, the results in FIGS. 2B and 2C were obtainedand are graphically represented in FIGS. 2D and 2E.

FIG. 2A shows an apparatus for observing the primary features of oneembodiment of the decarbonation portion of the present invention. Theapparatus in FIG. 2A (or one with similar effect) can be operatedaccording to the procedure listed below:

-   -   (1) Carbonation Unit 801 is charged with NaOH with a test load        (for example, 1M NaOH in 25° C. water) to a depth of 30 cm,        packed or unpacked.    -   (2) Flue-gas, simulated or actual, in this case, for a typical        coal-fired flue-gas exhaust (16% CO₂, 84% N₂, SO_(X)/NO_(X) <1%        or in natural ppm rates) is introduced to Carbonation Unit 801,        effectively sparged in an unpacked column, effectively sparged        or distributed in packed columns, travels through the fluid and,        is vented. Gas is at 25° C., CO₂ is 2L/min with other gases        mixed proportionally, flowing upwards through a circular 4″        diameter column; system pressures can be 1 atm psig or less.    -   (3) It can be noted by means of measuring incoming CO₂        concentration (by gas chromatograph sampling, for instance, or        with in-line measurement of CO₂ concentration) that CO₂ is being        absorbed by the fluid, that temperatures are rising (exothermic        reaction), and liquid assay sampling will show the presence of a        carbonate/bicarbonate/hydroxide equilibria, indicating that not        only absorption of CO₂ is occurring, but its conversion into        carbonate or bicarbonate form is proceeding. Practical operating        experience indicates that these key “transition points” exist in        the pH equilibria:        -   a. At pH<=8.3, the formation of bicarbonate is favored.        -   b. At pH>=10, the formation of carbonate is favored.    -   (4) The absorption/conversion to carbonate reaction proceeds        strongly and exothermically until, given the flow dynamics of        the gas, at whatever rate the CO₂ was being absorbed/converted,        the exothermic phase of the reaction ends, temperature plateaus        first and then falls, and the absorption capability of the        fluid, that falls as the OH ion concentration decreases,        effectively zeroes at this point. pH generally closely        approximates 8.3 or in its near-neighborhood when absorption        levels begin to fall; at pH>8.3, absorption is relatively        robust.    -   (5) The fluid is transferred to Bicarbonation Column 803, and        flue-gas is again introduced to the fluid. Absorption of CO₂ has        ceased and in some cases, will be shown to be negative (the        fluid gives up some CO₂ to the gas flow traveling through it).        The temperature of the fluid continues to fall, partially due to        some incidental evaporation to the migrating gas stream, but        also due to the reaction of bicarbonation that is taking place        between the previously created sodium carbonate and the        remaining “orphan” CO₂ that is dissolved in the fluid.    -   (6) The equilibrium continues to be shifted toward bicarbonate,        and optimizations of starting hydroxide concentration, fluid and        gas temperatures, pressures, flow-rates and velocities,        tortuosities, etc. can be accomplished, even up to the point of        producing pure bicarbonate (99%+).

FIGS. 2B and 2C depict the results of several test-series conducted witha charged-load (a specific concentration of NaOH was placed in adecarbonation system as shown in FIG. 2A). Several key points aredemonstrated by the data in FIGS. 2B and 2C:

-   -   (1) Conditions can be modified sufficient to reproducibly create        either pure carbonate (runs 4 and 14), or pure bicarbonate (runs        28 and 32), and may be modulated to achieve various results (or        “ion ratios” between the extremes of 1.0-2.0).    -   (2) The reactor dimensions that result from this study are, for        all cases with significant absorption, found to be that        gas-liquid contact distances generally less than 3 m are        sufficient to achieve 90% absorption of incoming gases. Hence,        short, low-resistance stages are shown to be designable so as to        achieve high rates of absorption consistent with the        thermodynamic efficiency limits. In other words, the physical        process of removing the CO₂ operates at absorption levels that        may apparently meet or exceed the thermodynamic efficiency of        the system. Such high absorption rates (gas in, gas out) do not        account for the energy and, therefore, CO₂-production.        Therefore, keeping CO₂ absorption rates (from the fluid) and        thermodynamic efficiency of the plant as two clearly different        measures is important to avoid confusion.

The results from FIGS. 2B and 2C (absorption of CO₂ by the fluid) andthe product ion ratio (1.0=bicarbonate, 2.0=carbonate) are depicted inFIGS. 2D and 2E. Several important conclusions can be derived from FIGS.2B and 2C:

-   -   (1) Instantaneous absorption-rates as high as 98% in a single        absorption-stage of incoming CO₂ are noted.        -   a. Pure bicarbonate (NaHCO₃) was produced in solution at            conditions that absorbed 25% of incoming CO₂ in a            single-stage bubble-column gas-liquid contactor with a depth            of 0.30 m fluid depth/gas-liquid contact distance.            Extrapolating to a 90% absorption, 3 meters of contact            distance is sufficient to absorb 90% of incoming CO₂.        -   b. Pure carbonate (Na₂CO₃) was produced in solution at            conditions that absorbed 70% of incoming CO₂ in a            single-stage bubble-column gas-liquid contactor with a depth            of 0.30 m fluid depth/gas-liquid contact distance.            Extrapolating to a 90% absorption, <2 meters of contact            distance is sufficient to absorb 90% of incoming CO₂.        -   c. Various absorption vs. carbonate ion ratios in products            indicate that a continuum of solutions exist between these            extremes.    -   (2) The absorbent fluid retains its absorption characteristics        for industrially-worthy lengths of time (e.g., 15-240 minutes in        these examples).    -   (3) The reactor input variables (concentration, temperature,        pressure, gas flow-rate, contact time, etc.) can be modulated to        produce pure bicarbonate, pure carbonate, or any mixture        in-between.    -   (4) Using these laboratory results to design a 90% CO₂ reactor        results in solutions under 3 m of gas-liquid contact distance        (e.g., approximately fluid depth, column height), and 1 m in        many industrial-worthy process corners.

Example 4 Analysis of LVE for Various Chemical Conditions

FIG. 5 is a chart showing low-voltage electrolysis operating lines forvarious chemical conditions. It depicts some typical experimentalresults, in which a membrane chlor-alkali cell is operated undernon-standard conditions, namely:

-   -   (1) pH of the anolyte fluid (protonated brine) is adjusted by        closed-loop pH controlled addition of HCl (hydrochloric acid in        water) at pH of 1.0, 2.5, and 5.0;    -   (2) temperature of the anolyte fluid is held at setpoint, by        closed-loop fluid circuits heated by electric heaters; and    -   (3) voltage is modulated for each fluid/protonation/temperature        condition, with the current attained by the 0.01 m² chlor-alkali        cell recorded.

In FIG. 5, note the set of example experimental data, which plots actualexperimental voltage vs. current (translatable into current densities,kA/m², as indicated on the chart) for a 0.01 m² electrolysis cell, 13 mmgap, operating at various combinations of temperature and degree ofprotonation of the anolyte brine fluid (controlled in this experimentalseries by closed loop ph-control of HCl (l) addition to the brine loop).

Note the following concerning these typical results in FIG. 5:

-   -   (1) At high-voltages (5V) such as are normally used in the        chlor-alkali use of such cells, the maximum current (and for a        given cell, therefore the maximum current density) is attained.    -   (2) Higher temperature brine at the same pH has a superior        current density at a given voltage.    -   (3) Lower pH brine has a superior current density at a given        voltage compared to higher pH brine.    -   (4) These general tendencies (higher temperature, higher acid        concentration) can be optimized by standard        design-of-experiments techniques for each individual        electrochemical cell geometric/component design to produce the        optimum (kA/m²V) for that cell. Similar experimentation on any        chlor-alkali cell with increased operating pressure will result        in concluding that increased operating pressure also enhances        (kA/m²V).    -   (5) The slope of the lines (ΔV/ΔA) is initially large, with        relatively large drops of voltage occurring with relatively low        drops in current/current-density; however, after an inflection        point is reached (at approximately (2.5V, 10 A/0.01 m²)),        further reductions in voltage result in more extreme reductions        in current and therefore in current density.    -   (6) This inflection-point and its near-neighborhood of operating        conditions represent the optimum voltage-vs-current-density        trade-off in an economic sense of efficiency. Standard        design-of-experiments optimization can achieve the optimum        low-voltage condition for any physical cell embodiment of the        invention.    -   (7) In the context of this example alone, that the 1.0 pH,        90° C. anolyte condition has a superior current/voltage        characteristic and is therefore the optimum operating line        represented among these various demonstrated operating lines.    -   (8) The primary drawback of low-voltage electrolysis is the        accompanying decrease in current-density; kA/m² declines with        declining voltage. Since the system must produce the same number        of Na+ ions to absorb the same amount of carbon dioxide, the m²        area of the membrane surfaces must proportionally increase;        e.g., if current density drops by 50%, then twice as much        membrane area will be required to produce sufficient absorbent        fluid. This has a serious effect on plant cost, as a        chlor-alkali plant has costs that are nearly proportional to        membrane area. Low-voltage electrolysis offers several        advantages that may, in certain embodiments of the invention,        allow optimization along low-voltage lines that redress        significantly or wholly for this large-area requirement        drawback. Namely, lifetimes of membranes and electrolytic cell        components that are operated in more benign/less energetic        operating conditions that can extend cell and/or membrane life        may be experienced. Designing specifically for lower-voltage        conditions may attend some ability to relax certain materials        and performance criteria that are not as essential in        embodiments that employ low-voltage. Certain of these        degrees-of-design freedom may result in low-cost cells that        partially or wholly absorb the incremental cell membrane cost        originally incurred due to low-voltage/low-current-density        operation. In these and many other ways, LVE systems, while        requiring larger membrane areas than standard chlor-alkali cells        for the production of the same amount of NaOH, may wholly or        partially assuage some of that additional cost and operational        expense.    -   (9) The trade-off between the benefit (lower voltage and hence        lower power) and the detriment (higher area and degrading        current density) may be optimized by the technique described in        Example 7. For the 1.0/90° C. operating line depicted in FIG. 5        (which for this small example set is the superior V/I        characteristic for LVE operation), there can be calculated a        Voptlve, and from the above relationship, the Ioptlve can be        obtained. Hence, for a given electrolysis cell geometric design,        conditions of temperature, pressure, brine concentration, degree        of protonation, membrane choice, etc. may all be done to produce        a superior V/I curve or operating line, and then the optimum        point on that curve can be calculated by the method of        Example 7. In this case, the Voptlve is 2.88V and the current        density Ioptlve is 1.04 kA/m².    -   (10) In Example 7, the current at Vopt=2.88V is approximately 5        A in the lightly-protonated and/or low-temperature cases. In        just this example, that current (and therefore current density)        was more than doubled to 10.4 A.    -   (11) Additional protonation of the brine, temperature, pressure,        concentration, geometric arrangement of the components of the        cell, electrical fields, and conditions can be similarly        optimized to produce a superior (kA/m²V) metric, but protonation        itself increases the amount of stoichiometric hydrogen produced,        thereby increasing the energy pay-back of the system. It is        important to note that optimizing for the lowest-energy CO₂        absorption/conversion can be attained by both optimizing the        (kA/m²V) of the system, which lowers the energy required to        manufacture the absorbent fluid, but that simultaneously        optimizing the hydrogen available for energy-recovery (and then        optimizing the efficiency by which that available hydrogen        energy is recovered), the entire energy for the process may be        optimized to its lowest potential, for a given physical        electrochemical cell of specific design.    -   (12) Given that embodiments of the present invention can        effectively absorb CO₂ in extremely dilute hydroxides (0.2M and        less have been demonstrated) compared to the concentrated        hydroxide normally produced in chlor-alkali manufacture        (typically 33-35% by weight, then concentrated further by        steam-evaporation), the design of chlor-alkali cells for        low-concentration operation (as well as low-voltage operation)        may provide new degrees of freedom for design optimization at        these non-standard conditions.

Embodiments of the present invention are incapable of making morehydrogen-energy than the energy consumed in making that hydrogen.Otherwise, a violation of the Second Law of Thermodynamics would result.This places a limit on the minimum voltage that can be applied to theelectrochemical cell. Presuming 100% efficiency on the hydrogen return,and using 39000 kw-hr/ton H₂ energy content (EIA reference value), wouldresult in a minimum voltage of 1.55V. A person of skill in the art can,for any system with a given hydrogen/electric return efficiency and achosen value for the energy content of the system, compute a minimumachievable voltage for that system.

In practice, thermodynamic inefficiencies (including but not limited to,I2R losses, current inefficiency in the cell, waste-heat losses, etc.),and the requirement of slight over-voltage to operate, raise the minimumvoltage attainable for a given cell. The above figures vary slightlydepending upon the value of “a,” the protonation ratio, as it varies theamount of hydrogen available for energy-recovery.

That said, current-density at low-voltage determines the amount ofelectrolysis area (a good scalar for capital expense) required toproduce an amount of caustic, and at minimum voltage, the area requiredis extremely large. Hence, some voltage above minimum voltage isrequired for operation, the amount depending upon thecapital-expense/ecological-efficiency trade-off chosen in the design.Current efficiency (the percent of current spent in manufacturingproduct) declines at low voltage, so optimizing low-voltage electrolysisoperation is not the same as attaining a single low-voltage operatingcondition. Current processes are designed to operate in the LVE regime(below 5V), and at these voltages below 5V power consumption overtraditional techniques is significantly enhanced.

Example 5 Thermodynamics of a Large-Scale Plant Design

For this example, a model plant (incorporating certain embodiments ofthe present invention) exhibiting full-scale operating plant behavior isexplained and the energy required to extract a given amount of CO₂quantified and bounded within statistical limits by various means andmethods, including the following:

-   -   (1) Thermodynamic efficiency (∂CO₂/∂E) can be approximated as        ΔCO₂/ΔE over sufficiently short range intervals of E (energy).    -   (2) Certain simplifying assumptions can be made regarding a        plant design such as is represented in FIG. 9A, among them:        -   a. The primary energy spent is in the electrolysis process;            pumping, compression, controls, etc. are considered de            minimis relative to the energies spent making reactants            (electrolysis) and in hydrogen energy-recovery. These values            assumed zero or <0.1% of power consumed in electrolysis            operations.        -   b. Electrolysis energies spent can be represented            approximately by the following equation:            Eout=V*I*EFFcurrent        -   Where:            -   V=voltage of operating electrolysis cell            -   I=current required to produce the chemicals by                electrochemical half-reaction, including the                greater-than-1:1 stoichiometry caused by the protonation                of the brine. 0.05 HCl/NaCl ratio of protonating ions                consumed in electrolysis is used in the example.            -   EFFcurrent=the current efficiency, defining the amount                of current used in the actual production of chemical                species, with the remainder being lost in I2R losses,                etc. 97% is the value used in the example; each                electrolysis cell will have its own unique current                efficiency, which degrades and varies over the life of                the cell.        -   c. Energy recovered from hydrogen-combustion (by whatever            means, combustion as a boiler gas, combustion in a            fuel-cell, etc.) as:            Ein=39000 kw-hr/toncompressed H2*Ton H2*EFFdc        -   Where:            -   Ton H2=tons of hydrogen produced by the process                including the hydrogen produced by the greater-than-1:1                stoichiometry, caused by the protonation of the brine.            -   EFFdc=the efficiency of the hydrogen-recovery process in                converting the incipient energy of the hydrogen gas into                DC electric current. 60% of the hydrogen-energy produced                is recovered into DC current. This is the current                off-the-shelf DC efficiency of hydrogen/atmospheric                oxygen fuel cells; a similar number is obtained for the                process of compressing the hydrogen (which costs 15% of                its power to accomplish) and then assuming 85% of the                hydrogen energy is returned in the customer processes                such as hydrogenation processes, gasoline reforming                processes, etc.        -   d. Energy returned from waste-heat recovery salvaged from            the heat of the incoming flue-gas stream. Incoming heated            gases are cooled once entering the process. In some            embodiments of the invention, this cooling can be            accomplished by absorption of the waste-heat and the            conversion of that heat into electrical DC energy, which can            be used to supplement/fully-power/over-power the process            that composes the invention. In this example, the            supplemental waste-heat recovery is not included.

The plant model for this example includes a modeling of the flue-gasexiting the power-plant under normal operating conditions, as shown inFIG. 9B. This involves significant assumptions regarding the compositionof the fuels, the efficiency of the combustion processes themselves, therelative proportion of elements in the combustion processes, etc. Theassumptions for this example are depicted in FIG. 9B and are consistentwith the flue-gas output of a typical sub-bituminous coal-fedpower-plant with a 10,000 BTU/kw-hr heat-rate.

For a given flue-gas output, there is a hydroxide requirement that canbe calculated. Several assumptions are required here. The ratio of ions(“ion ratio” is the ratio Na/C in the absorption/conversion reaction) isthe same as the ratio of those elements in the product solids formed. Inthe case of pure bicarbonate, the number would be 1.0, in the case ofpure carbonate, the number would be 2.0, and for mixtures of bicarbonateand carbonate, the number would lie between 1.0 and 2.0. The calculationfor the caustic requirement for this example is depicted in FIG. 9C. Forthe example depicted in FIG. 9C, the ion ratio is 1.0.

For a given hydroxide requirement, there is a corresponding electrolysisrequirement, based upon the amount of water, salt, square meters ofmembrane surface (the scalar for electrochemical cells of this kind),and current density (itself a function of the cell design, chemistry,and operating conditions; here the figure of 3 kA/m² is used). Thesecalculation of the electrolysis requirement for this example is depictedin FIG. 9D.

For a given amount of electrolysis under protonated conditions, there isa given amount of hydrogen gas produced, which itself represents acertain amount of energy available for recovery, or the hydrogen is usedchemically in further processing. In this example, the hydrogen isconverted to DC electricity using a fuel cell operating at EFFdc=60%.

For a given amount of flue-gas processed, there is a certain waste-heatcontent that can be extracted from it at a certain efficiency ofconversion into DC electricity, and that recovered electricity may beused to supplement the DC electricity consumed by the process inelectrolysis. The waste-heat for this example is depicted in FIG. 9E,with an efficiency chosen, in this case, of 25%, a figure that isexceeded by various waste-heat/DC generation techniques extant in thefield.

Given these individual components of energy inputs and outputs, the neteffect of these energy transfers may be summed, as in FIG. 9F. Here,energies are presented in kw-hrs and as percentages of the basis-plantpower, and a calculation of ecological efficiency for this example isshown.

In some embodiments of the present invention, an additional H₂/Cl₂ fuelcell may be employed to combust hydrogen and chlorine gases for thepurpose of recycling HCl used for protonation of the brine. Inparticular, the amount of “super-stoichiometric” HCl can be recycled,and in theory, eliminate the need for stock chemical HCl to be added tothe system. In actual practice, a certain amount of make-up HCl must beperiodically added to the system. The combustion of H₂ in Cl₂ containsmore energy than does the H₂/O₂ combustion. Slight additional energy isthereby attained. However, fuel-cells are inherently somewhatless-than-perfectly efficient, so the energy “gain” from using chlorineas the oxidant is more than outweighed by the loss inherent in therecovery method. Hence, in this example, these effects are considered tocancel each other out, and are zeroed. Therefore, the two effectssomewhat offset each other but still result in a net loss. However, theeffect of the presence of the additional protons in the electrolysis isto dramatically catalyze the production of NaOH at low voltages and high(kA/m²V) at those low voltages. Hence, for any given apparatus, anoptimization can be carried out to recycle a given amount of H₂/Cl₂ intoHCl and to protonate the incoming brine with that amount of HCl. At someoptimum value (usually found between a=0.05 and a=1.0M, or near pH=1 at90° C.), hydrogen/chlorine fuel cell losses (which outweigh the slightgain over oxygen oxidation presented by the chlorine) and hydroxideenergy benefits (better kA/m²V) will be simultaneously optimum for theentire system. It should be noted that in this example, only H₂/O₂combustion is calculated; the H₂/Cl₂ combustion has a slightthermodynamic gain from the extra strength of the chlorine oxidation,but the countering effect of fuel-cell inefficiency makes for a slightlynegative, but considered de minimis, effect.

A. Calculation of Ecological Efficiency

Calculation of ecological efficiency (∂CO₂/∂E) and ΔCO₂/ΔE for thisexample was accomplished as follows:

-   -   (1) It was presumed that there were three plants:        -   a. The basis power plant (exemplified in the flue-gas model            in FIG. 9B)        -   b. The CO₂ Absorption/Conversion plant (which requires            supplementary power to process the flue-gas from the basis            power plant and returns a portion of that power from            hydrogen combustion or the calculated power inherent in            hydrogen, if hydrogen is the end-product and is not            combusted).        -   c. A third, supplemental power-plant that provides the power            required by the CO₂ Absorption/Conversion plant. In this            example, the characteristics of this power plant were            assumed to be identical to the basis plant.    -   (2) The following aspects relating to the CO₂ and energy spent        in processing 100% of the basis plant were then calculated:        -   a. CO₂ from the basis plant (flue-gas model)        -   b. Energy produced by the basis plant        -   c. Net energy required by the CO₂ absorption/conversion            process        -   d. Net energy required by the supplemental power plant is            assumed identical to the Net energy required by the CO₂            absorption/conversion process.        -   e. CO₂ produced by the supplemental plant is assumed to be            proportional to the energy produced by the supplemental            plant, and with the same ΔCO₂/ΔE of the basis plant.    -   (3) The following results for the above calculations were        obtained:        -   a. Basis Plant—a 10,000 heat-rate plant producing 1 Gw            continuously for a 1-year basis produces 8.76 Bkw-hrs each            year and produces a basis of 7,446,068 tons of CO₂ per year,            averaging 1176 kw-hr/ton CO₂.        -   b. CO₂ Absorption/Conversion Plant—for this example            calculation (a=0.10, 2.1V operation, pure bicarbonate            produced, 15% of hydrogen energy consumed in compression,            pumping/compression costs and waste-heat recovery benefits            excluded), 3.648 BKw-hr are required to absorb/convert 100%            of the basis plant.        -   c. Supplemental Power Plant—the plant in this example            produces the power required by the CO₂ Absorption/Conversion            Plant, 3.648 Bkw-hr, and itself produces (by the 1176            Kw-hr/ton CO₂ figure from above) a total of 3,101,209 tons            of CO₂ that is presumed to be emitted to the atmosphere.        -   d. Total power generated is therefore 12.48 Gw-hrs. Total            delivered power is therefore 8.76 Gw-hrs. Total CO₂            generated is therefore 10.55 Mtons. Total CO₂ emitted is            therefore 3.101 Mtons. 29.1% of total power is consumed in            the CO₂ absorption/conversion process. 71.9% of total CO₂ is            consumed.

Several key points are illustrated by the above calculations:

-   -   (1) Arithmetically, it is demonstrated that the following        formulae apply:        % power consumed=1−% CO₂ consumed        % CO₂ consumed=1−% power consumed    -   This forms a line, called the One Unit Operating line, shown in        FIG. 9A.    -   (2) For this example, the (∂CO₂/∂E) and ΔCO₂/ΔE are        algebraically identical, namely:        Δ CO₂/ΔE=(∂CO₂/∂E)=0.291/0.719=0.41

Further extrapolated cases can be further modeled, in which cases, theCO₂ emitted by the Supplemental Power Plant is itself treated by anotherCO₂ absorption/conversion process unit #2 of correspondingly smallercapacity, and that absorption/conversion unit #2 is correspondinglypowered by a Supplemental Power Plant #2, etc., which results in resultslike those in Table 2 for the first five series of iterations.

TABLE 2 Produces Power CO₂ Produces Amount Inc Power Plant (Mton) Power(kw-hr) Absorbed Required Basis 7,446,068 8,760,000,000 7,446,068  3,648,448,267    Iteration 2 3101209 1519540497  Iteration 3 1291621632872705 Iteration 4  537947 263584854 Ecological Power SupplementalEfficiency Plant CO₂ % tot E % tot CO₂ (∂CO₂/∂E) Basis 3101209 29% 71%2.40 Iteration 2 1291621 37% 89% 2.40 Iteration 3 537947 40% 96% 2.40Iteration 4 224049 41% 98% 2.40

Several points regarding Table 2 are significant to the model:

-   -   (1) Note that the efficiency of the process, whether in the        basis case, or any of the successive iterated cases,        consistently produces the same (∂CO₂/∂E) value for the system;        this term is considered constant for systems that approximate        the constraints of this model and is called, for these purposes,        the ecological efficiency of the process.    -   (2) It is clear that the value (∂CO₂/∂E) is constant in all        solutions, so a solution can be derived when then number of        iterations is presumed infinite; i.e., when the plant is        operated so as to consume 100% of the CO₂ produced by the plant,        by the simple expedient of using the following equation:        1/(∂CO₂ /∂E)=% of plant power required to absorb/capture 100% of        produced CO₂        -   In the example case, this calculates as 41.6%.    -   (3) Alternatively, it is evident that, when the net power spent        in absorption/conversion is zero (neglecting waste-heat        recovery), for a given process condition, the CO₂ absorbed and        converted is likewise zero. Hence, all operating lines for        plants of this type theoretically intersect at (0% Power, 0%        CO₂).    -   (4) Given any two points in a linear system, straight-line        solutions for Operating Lines may be constructed that define the        operating characteristics of the CO₂ Absorption/Conversion        process, by the following means:        -   a. for each operating condition, a basis-case solution is            accomplished, and the resulting point One Unit Case solution            (% Power, % CO₂) is plotted on a graph of % CO₂ (y-axis) vs.            % Net Power consumed α-axis);        -   b. for that case, (∂CO₂/∂E) is calculated, and the case at            y=100% is solved for the x-coordinate; and        -   c. all lines are presumed to travel through the origin. In            actual systems, there would be some power consumption            (controls, environment, etc.) at even zero absorption, so            this is an idealized case. In practice, these lines would be            slightly curved and not terminate at the origin.    -   (5) In this way, a family of operating lines for CO₂        absorption/conversion processes of this type can be created.    -   (6) On this same type of plot, competing technologies can also        be plotted and compared graphically, e.g.:        -   a. A competing MEA (methyl-ethyl-amine) absorption            technology consumes 30% of plant power to accomplish an            absorption of 58% of CO₂ emitted before absorption was            introduced.        -   b. Further, an estimated 15% of plant power is expended in            liquefying this CO₂ through extreme pressure and            refrigeration cycling (45% power/58% CO₂).        -   c. This would then demonstrate a (∂CO₂/∂E) value of 1.24;            however, there is additional unaccounted energy required to            transport/inject/maintain the CO₂ in a sequestration-site.        -   d. Graphically, this competing technology is shown to be            less efficient than the example CO₂ absorption/conversion            plant operating the process which is one embodiment of the            invention; i.e., this model shows the competing technology            would require 70%+ of the power plant to eliminate 100% of            its CO₂ production. Note these points regarding the            competing technology as graphically represented in FIG. 9A            (refer to legend on chart):            -   i. According to 2005 EIA estimates, the absorption of                CO₂ by a MEA technology requires 30% of the plant power                to absorb 58% of the flue-gas CO₂ produced. (Note                position on chart in FIG. 9A (30%, 58%) for                absorption-processing alone.)            -   ii. By the same estimates, compression/liquification of                that CO₂ consumes another 15% of the plant power, moving                the operating point of such a plant to (45%, 58%).            -   iii. There is no firm estimate of the energy required to                transport the liquid CO₂ by pipeline or other                transportation device, nor for that matter the amount of                energy necessary to pump or inject that CO₂ into a                carbon store of various natures, nor the amount of                energy that might be required to maintain that CO₂ in                said stores for perpetuity. However, though those                additional energies are not estimable, it seems                reasonable to assume they are non-zero. Hence, the                ecological efficiency of such a device is logically                worse than a (45%/58%) trade-off in power-spent to                secure a certain CO₂-reduction benefit. Extrapolating                this to a 100% remission case, the                MEA/liquification/sequestration technique would consume                more than 70% of the plant power. It should be noted                that typical competing absorption technologies cannot                approach 100% absorption; i.e., the figure of 58% CO₂                absorption was for a plant that processed 100% of                outgoing flue-gas.

B. Calculation of the Limits of Ecological Efficiency [(∂CO2/∂E)max]

In practice, for a given system that effectively converts all the NaOHproduced to NaHCO₃ by absorbing CO₂, the primary energy component is thekw-hr/mole NaOH. Although the power per mole NaOH is proportional toboth voltage and current, the current is fixed by the stoichiometry ofthe chemistry. Thus, the power expended per mole CO₂ is primarilyoptimized by achieving the lowest voltage condition that efficientlyproduces hydroxide.

The minimum voltage at which an electrolysis system according toembodiments of the present invention operates (as configured withvarying concentrations, geometric dimensions, flow-rates, etc.) can bedetermined by observing the Current-Density (kA/m²) vs. Vcharacteristics of the system and determining the lowest voltage atwhich sufficient, non-zero current densities are obtained to makeproduct. Altering the physical dimensions, electrical field generationdevices, cell geometries, compositions of materials, and processingconditions to optimize this characteristic metric (kA²/m²V) is a primarymeans to optimize these systems, and typical design-of-experimentstechniques are useful for optimizing an industrial process for a givenphysical plant.

Practical limitations aside, there is one fundamental limit that willapply to all systems with a given H/Na ratio (protonation ratio),namely:

-   -   (1) No device can operate that produces more energy through        hydrogen-energy recovery than is input to the system in        electrolysis. Persons familiar with thermodynamic principles        will note this would be a “Second Law Violation.”    -   (2) As a result of this fact, a fundamental thermodynamic limit        can be bounded, given a choice of H/Na ratio used in the anolyte        consumption:        -   a. For this example, H/Na was presumed to be 0.10.        -   b. The hydrogen energy return efficiency was set at 100%.        -   c. The lowest voltage at which operation can occur, in which            the net energy consumed by the system is zero (“Vmintheo”)            (i.e., the point where electrolysis costs equal the assumed            100% hydrogen return efficiency), was calculated.        -   d. In this example, that low voltage is 1.5526V. This number            is a strong function of the Na/C ratio, the H/C ratio, and            the hydrogen-energy return efficiency. In this optimal case,            Na/C is 1.0 and H/C is 1.0.        -   e. Following this calculation through to its ecological            efficiency, the Single Unit solution is approximately 7%            power for 93% CO₂ absorption/conversion.        -   f. Processing at more efficient operating points than this            theoretical minimum is possible by:            -   i. supplementing the power consumption with waste-heat                recovery; and            -   ii. powering the absorption/conversion process either                partially or wholly with power whose production does not                cause CO₂ emission (hydro-electric, solar, wind,                nuclear, etc.).

Similarly, presuming ideal hydrogen-return efficiency, etc. as above,the maximum voltage at which operation can be “ecological” (“Vmaxeco”)(i.e., in which the CO₂ absorption/conversion process removes more CO₂than it creates) was calculated:

-   -   a. H/Na, Na/C, and hydrogen return energy efficiency were set at        1.0, 1.0, and 100%, respectively, as above.    -   b. The voltage at which the CO₂ removal would be 50% was        calculated.    -   c. In this example, that Vmaxeco is 4.457V. At this voltage and        condition, the process operates on the line ∂CO₂=∂E, the        boundary between ecologically beneficial and ecologically        harmful operation.

Hence, ecologically-beneficial operation will occur when theelectrolysis system is operated between Vmintheo (1.5526V) and Vmaxeco(4.457V). Operation between those two points may be replicated with manytypical electrolysis systems. Laboratory results at or below 2.1V may bereadily reproduced by manipulation of geometry, concentration,temperature, pressure, flow-rate, etc. of electrochemical cells designedin this fashion.

C. The Effect of Non-Greenhouse-Generating Power on EcologicalEfficiency

Where the supplemental power (that which powers the process) is producedby non-greenhouse-gas (GHG) emitting power (e.g., wind-power,hydro-electric, solar, nuclear, etc.), then there are zero supplementalCO₂ emissions, and the ecological efficiency of the present invention isvastly improved. For this example, the term 3101209 tons of CO₂ in Table2 is eliminated along with all the subsequent iterations, etc., leavingthis simplified result: all CO₂ is absorbed/converted (7,446,069 tons),and the total power required is simply the 8,760,000,000 basis plus the3,648,448,267 kw-hrs required to accomplish the work necessary toabsorb/convert that basis amount of CO₂, requiring only 29% of totalpower to secure 100% of the CO₂ emissions in the non-GHG-poweredprocess, compared to 41% of total power to secure 100% of the CO₂emissions in the GHG-powered process. This means that embodiments of thepresent invention provide a significant “leverage” factor when theprocess is powered by non-GHG emissions. Rather than using non-GHG powerto displace GHG-generating power on a 1%:1% basis, if the non-GHG poweris instead used to power processes that are some embodiments of thepresent invention, 1% of non-GHG-generating power then displacesGHG-generation by a multiplied factor, even in excess of theGHG-generating power examples described herein. One can easily envisioncases in which, for a given nation, state, or entity, a certainproportion of non-GHG-generating power, when used in this magnifiedfashion, could more efficiently attain any CO₂ reduction goal; i.e., onecould use “clean” power in a highly leveraged manner to clean-up theemissions of other “dirty power.”

Given that in some applications, non-GHG-producing power generation isavailable, sometimes in sporadic forms (e.g., solar, wind-power “farms”,etc.), the ability to utilize that power to make large quantities ofabsorbent during off-peak periods is extremely advantageous.

Example 6 Ecological Efficiency of Various Modeled Power Plants

FIG. 10 shows the ecological efficiency of various modeled power plantsincorporating embodiments of the present invention, and it depictsvarious conditions that are the primary factors in determining theecological efficiency, (∂CO₂/∂E).

From these calculations it can be generally concluded that:

-   -   (1) Forming sodium carbonate and using standard chlor-alkali        conditions, the process would have an ecological efficiency >1,        and while such an operation might be economically viable, it        produces more CO₂ than it absorbs.    -   (2) Altering the product equilibrium to favor the production of        sodium bicarbonate improves the ecological efficiency of the        process. In the case of altering conditions so as to produce        virtually pure sodium bicarbonate this advantage is        fully-optimized.    -   (3) Adopting low-voltage electrolysis practice firmly moves the        process into an operating region characterized by ecological        efficiencies lower than 1.0 (i.e., ecologically beneficial CO₂        absorption and conversion processes). Optimizing each physical        emulation of the electrolysis system for optimum (kA/m²V) and        maximum hydrogen energy production leads to further improvement        in ecological efficiency.    -   (4) Coupling the absorption/conversion process of embodiments of        the present invention with any number of available or        created-for-purpose machines that convert waste-thermal heat to        DC electricity, the initial energy investment in DC electrolysis        and AC pumping, etc. may be supplemented or wholly supplied from        the waste-heat recovery.

It should be noted that supplying the invention with power fromnon-green-house gas emitting power generators allows the process todirectly approach 100% CO₂ absorption (see discussion in Example 5).

Example 7 Determining Voptlve (the Optimum Low-Voltage Operating Voltagewith Respect to Cell-Capacity or Area) and Ioptlve (the Current at thatOperating Voltage), Given a V/I Characteristic Operating Line for anOptimized LVE Chlor-Alkali Cell

It has been demonstrated herein that lower-voltage operation lowers thepower required to manufacture sodium hydroxide that is used as anabsorbent fluid. Table 3 shows calculations made from the VI operatingline of the 1.0/90° C. anolyte case in FIG. 5 (discussed above inExample 4.).

Several points should be noted regarding the contents of Table 3:

-   -   (1) The third column, current efficiency (dimensionless),        represents the proportion of current generated that is used in        producing product chemicals; losses, such as I2R losses and        waste heating of electrolytic fluids are the primary cause of        inefficiencies. Current efficiency declines with declining        voltage.    -   (2) Cell area is normalized for a 3.975V case (where the current        density, and therefore the area requirement of the process, is        identical to a std 5V electrolysis running with 3.0 kA/m²        characteristic). A2/A1 (dimensionless) is calculated.

The last term, % of power-saved per dimensionless area, is plotted inFIG. 11. For such a function, the point at which the maximum slope(change in power usage per change in voltage) represents an optimum;i.e., at low voltage (e.g. 2.1-2.5V) the slope (Δpower/Δm²) isrelatively low, then at higher voltages (e.g. 2.5 to approximately 3),the slope (Δpower/Δm²) increases to a larger value, and then declines tolower slopes as the voltage continues to increase. This illustrates thatthere is a region of high-slope bounded by a region of low-slope on eachside; i.e., on either side of that Voptlve point, the change in powerusage per voltage delta is less effective.

A function that closely approximates the actual behavior wasaccomplished first (note the formula of the polynomial trend-lineproduced by least-squares fit). In this example:y=−10.164x3+88.256x2−235.06x+198.37 is a close-approximation. Then, thefirst-derivative of the function was calculated by typical treatment ofpolynomials: dy/dx=(3)(−10.164)x2+(2)(88.256)x−235.06=max. Values of x(Volts) can be iterated to find the maximum of this first derivative,which can be accomplished by various techniques, resulting in 2.894V asthe solution.

Note that voltages lower than 2.894V may be employed, and further powersavings will result. Some preferred embodiments will optimize low-poweroperation below this “natural optimum point.” In those cases, theadditional area used in the membrane will result in a “sub-optimized”electrolysis system, but low-power operation for the overalldecarbonation process may be further benefited by operating below thisnatural optimum for a given electrolysis sub-system. However, when doingso, the voltage/power benefit is thereafter attenuated, while thearea-factor continues to proportionally make operation less efficientper area.

The current and current density that corresponds to this Voptlve can bedetermined by either forming a similar least-squares relationship for Vand I, or by graphically using the operating curve to determine Ioptlve.In this example, the calculated value is 10.419 A (or for a 0.01 m² cellarea, as in this case), 1.042 kA/m².

TABLE 3 Volts, V operating conditions: I, Anolyte pH Power = Current1.0, 90 Current Δ(kA/m²)/ ΔPower/Δ(kA/ V * I * (A) Deg. C. Efficiency ΔV m²) Eff 30 3.975 97% 1.82 0.55 115.67 25 3.7 96% 1.87 0.53 88.83 203.433 94% 1.87 0.54 64.62 15 3.165 91% 1.80 0.56 43.34 10 2.887 88% 1.790.56 25.30 7.5 2.747 83% 1.74 0.58 17.15 5 2.603 78% 1.69 0.59 10.19 42.544 73% 1.49 0.67 7.41 3 2.477 67% 1.43 0.70 4.98 2 2.407 60% 1.190.84 2.90 1.5 2.365 54% 0.89 1.12 1.92 1 2.309 49% 0.64 1.57 1.13 0.72.262 44% 0.61 1.65 0.70 0.5 2.229 40% 0.62 1.60 0.44 0.4 2.213 36% 0.323.10 0.31 0.3 2.182 32% 0.56 1.80 0.21 0.2 2.164 29% 0.23 4.40 0.12 0.12.12 26% 0.00 212.00 0.05 I, % of STD % Power Current Cell Area Power/Saved by Δpowersaved/ (A) A2/A1 kgmole Reduced V m² 30 1.000 77% 23% 251.302 71% 29% 22.22 20 1.790 65% 35% 19.77 15 2.669 58% 42% 15.82 104.572 51% 49% 10.80 7.5 6.744 46% 54% 8.05 5 11.357 41% 59% 5.22 415.619 37% 63% 4.03 3 23.248 33% 67% 2.87 2 39.873 29% 71% 1.78 1.560.121 26% 74% 1.24 1 102.631 23% 77% 0.75 0.7 166.292 20% 80% 0.48 0.5262.506 18% 82% 0.31 0.4 367.228 16% 84% 0.23 0.3 551.770 14% 86% 0.160.2 927.267 12% 88% 0.09 0.1 2103.359 11% 89% 0.04

All of the methods and devices disclosed and claimed herein can be madeand executed without undue experimentation in light of the presentdisclosure. While the methods and devices of this invention have beendescribed in terms of preferred embodiments, it will be apparent tothose of skill in the art that variations may be applied to the methodsand devices and in the steps or in the sequence of steps of the methoddescribed herein without departing from the concept, spirit, and scopeof the invention. More specifically, it will be apparent that certaincompositions which are chemically related may be substituted for thecompositions described herein while the same or similar results would beachieved. All such similar substitutes and modifications apparent tothose skilled in the art are deemed to be within the spirit, scope, andconcept of the invention as defined by the appended claims.

REFERENCES

The following references, to the extent that they provide exemplaryprocedural or other details supplementary to those set forth herein, arespecifically incorporated herein by reference.

-   “Annual Energy Outlook 2005 (AEO2005),” prepared by the Energy    Information Administration, available through National Energy    Information Center, EI-30, Washington, D.C.-   “Carbonate Chemistry for Sequestering Fossil Carbon,” by Klaus S.    Lackner, in Annual Review of Energy Environment, 2002, by Annual    reviews.-   “Effects of the Operating Pressure On The Performance of    Electrolytic Cell at Elevated Temperatures,” by Y. Ogata, M. Yasuda,    and F. Hine, Nagoya Institute of Technology, Japan. In Proceedings    of the Symposium on Electrochemical Engineering in the Chlor-Alkali    and Chlorate Industries, The Electrochemical Society, 1988.-   “Electrochemical Hydrogen Technologies—Electrochemical Production    and Combustion of Hydrogen,” edited by Hartmut Wendt, Institute fur    Chemische Technologie, Stuttgart, Elsevier Press 1990.-   “Electrochemical Process Engineering—A Guide to the Design of    Electrolytic Plant,” by F. Goodridge and K. Scott, University of    Newcastle upon Tyne, 1995.-   “Exergy Study of the Kalina Cycle,” by Goran Wall, Chia-Chin Chuang,    and Masaru Ishida, presented at 1989 American Society of Mechanical    Engineers (ASME), Winter Annual Meeting, San Francisco, Calif.,    December 1989, published in “Analysis and Design of Energy Systems:    Analysis of Industrial Processes AES Vol. 10-3, pp 73-77 ASME.-   “Industrial Electrochemistry,” 2^(nd) Edition, edited by Derek    Pletcher and Frank Walsh, 1990.-   “Modern Chlor-Alkali Technology,” edited by M. O. Coultier, Society    for the Chemical Industry, London, 1980.-   “Modern Chlor-Alkali Technology,” Volumes 1-7, The Royal Society of    Chemistry Information Services, 1998.-   “Some Guidelines for Analysis, Design, and Operational Principles of    Bubble Column Reactors,” and other selected information contained in    “Bubble Column Reactors,” by Wolf-Dieter Deckwer, Gesellschaft fur    Biotechnologische Forschung mbH, Braunsweig, Germany, translated by    Robert Field, 1991 ISBN 0-471-91811-3.-   “Transport and Structure in Fuel Cell Proton Exchange Membranes,” by    Michael Anthony Hickner, dissertation submitted to the faculty of    Virginia Polytechnic Institute and State University, 2003.-   FIG. 8. “CO2 Absorption vs. (Low) Interfacial Gas Velocity at    Various Column Heights of 0.18M Sodium Hydroxide Solution in an    Unpacked Sparged Bubble Column,” from A. Schumpe thesis, University    of Hanover, 1978.-   Klara in: EIA Emissions of Greenhouse Gases in the U.S. 2000, EPGA    3^(rd) Annual Power Generation Conference, Hershey Pa., 2002.-   Mandal, et al., J. Chem. Engineering (Canada), 81:212-219, 2003.-   Shah et al., AiCHE J., 28(3):353-379, 1982.-   Unit Operations of Chemical Engineering, McGraw-Hill, 3^(rd) edition    © 1976, “Gas Absorption” pp. 707-743, after Eckert.-   Wie-rong et al., J. Zhejiang University Science, ISSN 1009-3095,    2004.

1. A method of sequestering carbon dioxide produced by a source andcontained in a gas stream, comprising the steps of: (a) obtaining ahydroxide in an aqueous mixture; (b) admixing the hydroxide with carbondioxide produced by the source to produce bicarbonate products or acombination of carbonate and bicarbonate products in an admixture; and(c) separating said products from the admixture, thereby sequesteringthe carbon dioxide in a mineral product form; wherein obtaining thehydroxide comprises: obtaining a group-1 or group-2 salt; admixing thesalt with acid and water, acid and steam, or acid, water, and steam toproduce a protonated brine solution; and electrolyzing the protonatedbrine solution to produce a hydroxide; wherein step (b) comprisesaltering the product equilibrium to favor the production of bicarbonateproducts.
 2. The method of claim 1, wherein an amount of carbon dioxideis generated in performing steps (a)-(c), and that amount of carbondioxide is less than the amount of carbon dioxide sequestered inperforming steps (a)-(c).
 3. The method of claim 1, wherein the admixingoccurs in two separate chambers, with one chamber being used to producecarbonate products and the other chamber being used to producebicarbonate products.
 4. The method of claim 1, wherein the admixingoccurs in a bubble column or series of bubble columns.
 5. The method ofclaim 1, wherein step (c) involves a heated-precipitation separationprocess.
 6. The method of claim 5, wherein the heat for the separationprocess is derived from heat exchange with incoming flue-gases.
 7. Themethod of claim 1, further comprising: transporting the products to aremote sequestration site; combining the products with acid in aneutralization reaction to generate pure carbon dioxide; and injectingthe carbon dioxide into a carbon bank.
 8. The method of claim 1, whereinother components of the gas stream are neutralized and/or captured inperforming steps (a)-(c).
 9. The method of claim 1, whereinelectrolyzing the protonated brine comprises combusting a carbon-basedfuel source to produce DC electricity.
 10. The method of claim 1,wherein the amount of acid admixed with the salt is based on adetermination of the optimum protonation rate that achieves the lowestenergy to produce reactants and the highest energy to recover fromproducts.
 11. The method of claim 1, wherein: the protonated brinesolution is electrolyzed in an electrochemical cell having a catholyteside and an anolyte side; and the products are recycled to the catholyteside of the electrochemical cell.
 12. The method of claim 1, wherein theenergy required by the method is supplemented with waste-heat recoveredfrom the gas stream.
 13. The method of claim 1, wherein the salt issodium chloride, the acid is hydrochloric acid, and the hydroxideproduced by electrolyzing the protonated brine solution is sodiumhydroxide.
 14. The method of claim 1, wherein the gas stream is anexhaust stream from a plant.
 15. The method of claim 14, wherein theplant is a power plant that employs a carbon-based fuel source.
 16. Themethod of claim 13, wherein electrolyzing the protonated brine solutionalso produces chlorine gas and hydrogen gas.
 17. The method of claim 16,wherein the chlorine gas and hydrogen gas are combusted to formhydrochloric acid.
 18. The method of claim 16, wherein the hydrogen gasis used to produce energy.
 19. The method of claim 17, wherein thehydrochloric acid formed from the chlorine gas and hydrogen gas is usedin obtaining additional hydroxide.
 20. The method of claim 16, whereinthe hydrogen gas is combusted with atmospheric oxygen or oxygen fromstock chemicals to produce water.
 21. The method of claim 18, whereinstep (c) involves a heated-precipitation separation process in which theheat for the separation process is derived from the energy produced bythe hydrogen gas.
 22. The method of claim 18, wherein the hydrogen gasis co-burned with coal to improve coal-fired emissions.
 23. The methodof claim 18, wherein the hydrogen gas used to produce energy is used ina combustion process for fuel-cell recovery of DC electricity.